Method and system for regenerating catalyst from a plurality of hydrocarbon conversion apparatuses

ABSTRACT

The present invention is directed to a method and system for integrating a catalyst regeneration system with a plurality of hydrocarbon conversion apparatuses, preferably, a plurality of multiple riser reactor units. One embodiment of the present invention is a reactor system including a plurality of reactor units, at least one reactor unit preferably comprising a plurality of riser reactors. The system also includes a regenerator for converting an at least partially deactivated catalyst to a regenerated catalyst. A first conduit system transfers the at least partially deactivated catalyst from the reactor units to the regenerator, and a second conduit system transfers regenerating catalysts from the regenerator to the plurality of reactor units. Optionally, catalysts from a plurality of hydrocarbon conversion apparatuses may be directed to a single stripping unit and/or a single regeneration unit.

FIELD OF THE INVENTION

[0001] The present invention relates to a method and system useful inhydrocarbon conversion processes and particularly in oxygenate to olefinconversion reactions. More particularly, the method and system of thepresent invention is to a plurality of reaction units, each reactionunit preferably containing multiple riser reactors, and an integratedcatalyst regeneration system.

BACKGROUND OF THE INVENTION

[0002] Light olefins, defined herein as ethylene and propylene, serve asfeeds for the production of numerous chemicals. Olefins traditionallyare produced by petroleum cracking. Because of the limited supply and/orthe high cost of petroleum sources, the cost of producing olefins frompetroleum sources has increased steadily.

[0003] Alternative feed stocks for the production of light olefins areoxygenates, such as alcohols, particularly methanol, dimethyl ether, andethanol. Alcohols may be produced by fermentation, or from synthesis gasderived from natural gas, petroleum liquids, carbonaceous materials,including coal, recycled plastics, municipal wastes, or any organicmaterial. Because of the wide variety of sources, alcohol, alcoholderivatives, and other oxygenates have promise as an economical,non-petroleum source for olefin production.

[0004] The catalysts used to promote the conversion of oxygenates toolefins are molecular sieve catalysts. Because ethylene and propyleneare the most sought after products of such a reaction, research hasfocused on what catalysts are most selective to ethylene and/orpropylene, and on methods for increasing the life and selectivity of thecatalysts to ethylene and/or propylene.

[0005] The conversion of oxygenates to olefins in a hydrocarbonconversion apparatus (HCA) generates and deposits carbonaceous material(coke) on the molecular sieve catalysts used to catalyze the conversionprocess. Excessive accumulation of these carbonaceous deposits willinterfere with the catalyst's ability to promote the reaction. In orderto avoid unwanted build-up of coke on molecular sieve catalysts, theoxygenate to olefin process incorporates a second step comprisingcatalyst regeneration. During regeneration, the coke is at leastpartially removed from the catalyst by combustion with oxygen, whichrestores the catalytic activity of the catalyst. The regeneratedcatalyst then may be reused to catalyze the conversion of oxygenates toolefins.

[0006] Typically, oxygenate to olefin conversion and regeneration areconducted in separate vessels. The coked catalyst is continuouslywithdrawn from the reaction vessel used for conversion to a regenerationvessel and regenerated catalyst is continuously withdrawn from theregeneration vessel and returned to the reaction vessel for conversion.

[0007] Conventionally, in order to produce an increased volume ofdesired product or to form different products, multiple, complete andindependent reactor systems with independent separation vessels wererequired. Each reactor in the multiple, complete and independent reactorsystems had its own regeneration system and/or stripping system. Withmultiple regeneration and/or stripping systems comes an attendantmultiplication of costs.

[0008] It is therefore desirable to reduce number of regeneration unitsand/or stripping units in order to reduce the tremendous costsassociated with implementing a plurality of multiple, complete andindependent reactor systems.

SUMMARY OF THE INVENTION

[0009] This invention provides a method and integrated multiple reactorsystem for converting a hydrocarbon over a catalyst to one or moreproducts while reducing the number of regeneration units and/orstripping units implemented therein. By reducing the number ofregeneration units and/or stripping units, the costs associated withmanufacturing and operating the multiple reactor system can besignificantly reduced.

[0010] In one embodiment, the invention provides a reactor systemincluding a plurality of hydrocarbon conversion apparatuses (HCA's),e.g., reactors or reaction units, and a regenerator for converting an atleast partially deactivated catalyst to a regenerated catalyst. Thereactor system also includes a first conduit system for transferring theat least partially deactivated catalyst from the HCA's to theregenerator, and a second conduit system for transferring theregenerated catalyst from the regenerator back to the plurality ofHCA's. The first conduit system optionally includes a first strippingunit for stripping the at least partially deactivated catalyst with afirst stripping medium. Also, the first conduit system optionallyincludes a second stripping unit for stripping the at least partiallydeactivated catalyst with a second stripping medium. The secondstripping medium can be the same as or different from the firststripping medium. Ideally, the first and second stripping units strip atleast partially deactivated catalysts from separate HCA's. One or moreof the HCA's preferably are multiple-riser reactors.

[0011] In another embodiment, the invention provides a reactor systemcomprising a first reaction unit comprising a first plurality of riserreactors, and a second reaction unit comprising a second plurality ofriser reactors, wherein each of the first and second reaction units hasa first end into which a catalyst can be fed and a second end throughwhich the catalyst can exit the reaction unit. The reactor system alsoincludes a regeneration unit having a regeneration inlet and aregeneration outlet, and a regeneration line having a plurality of firstline ends in fluid communication with the second ends of the first andsecond reaction units and a second line end extending to theregeneration inlet. A return line is also provided having a first returnend in fluid communication with the regeneration outlet, a second returnend directing a first portion of the catalyst to the first reactionunit, and a third return end directing a second portion of the catalystto the second reaction unit.

[0012] An alternative embodiment of the invention is a method forforming olefins in a methanol to olefin reactor system. The methodincludes contacting in a first reaction unit a first oxygenate with afirst catalyst under conditions effective to convert at least a portionof the first oxygenate to a first olefin and at least partiallydeactivating the first catalyst to form a deactivated first catalyst.The method also includes contacting in a second reaction unit a secondoxygenate with a second catalyst under conditions effective to convertat least a portion of the second oxygenate to a second olefin and atleast partially deactivating the second catalyst to form a deactivatedsecond catalyst. The deactivated first catalyst and deactivated secondcatalyst are directed to a regeneration unit and are regenerated to formregenerated catalysts. A first portion of the regenerated catalysts isdirected to the first reaction unit, and a second portion of theregenerated catalysts is directed to the second reaction unit. Thedeactivated first catalyst optionally contacts a first stripping mediumin a first stripping unit under conditions effective to removeinterstitial hydrocarbons from the deactivated first catalyst. Also, thedeactivated second catalyst optionally contacts a second strippingmedium in a second stripping unit under conditions effective to removeinterstitial hydrocarbons from the deactivated second catalyst.Alternatively, the deactivated second catalyst optionally contacts thefirst stripping medium in the first stripping unit under conditionseffective to remove interstitial hydrocarbons from the deactivatedsecond catalyst.

[0013] Another embodiment of the invention provides a hydrocarbonconversion system comprising first and second pluralities of riserreactors, each of the riser reactors having a first end into which acatalyst can be fed and a second end through which the catalyst can exitthe riser reactor. The hydrocarbon conversion system includes first andsecond catalyst retention zones provided to contain catalyst which canbe fed to the first and second plurality of riser reactors,respectively. Additionally, the system includes first and secondseparation zones into which the second ends of the first and secondpluralities of riser reactors, respectively, discharge the catalyst andproducts of a reaction conducted in the riser reactors. The separationzones are provided to separate the catalyst from the products of thereaction conducted in the first and second pluralities of riserreactors. First and second catalyst returns are also provided in fluidcommunication with the first and second separation zones, respectively,and the first and second catalyst retention zones, respectively. Thehydrocarbon conversion system also includes a regenerator forregenerating the catalyst, and first and second catalyst outlet lines,each of the outlet lines having a first end into which a catalyst can befed from the first and second pluralities of riser reactors,respectively, and a second end through which the catalyst can enter theregenerator. The system also provides first and second catalyst returnlines, each of the return lines having a first end into which a catalystcan be fed from the regenerator and a second end through which thecatalyst can enter the first and second pluralities of riser reactors,respectively.

[0014] Yet another embodiment of the invention provides a catalystregenerator system comprising a regeneration zone for contacting an atleast partially deactivated catalyst with a regeneration medium underconditions effective to form a regenerated catalyst, a plurality ofcatalyst inlets for receiving the at least partially deactivatedcatalyst from a plurality of reactor units, and a plurality of catalystoutlets for delivering the regenerated catalyst to the plurality ofreactor units. A stripping zone or zones optionally is provided forcontacting the at least partially deactivated catalyst with a strippingmedium under conditions effective to remove interstitial hydrocarbonsfrom the deactivated catalyst.

[0015] In one embodiment, the invention is a method for regeneratingcatalyst comprising receiving an at least partially deactivated catalystfrom a plurality of multiple riser reaction units, heating the catalystunder conditions effective to convert the at least partially deactivatedcatalyst to a regenerated catalyst, and directing the regeneratedcatalyst to the plurality of multiple riser reaction units.

[0016] Another embodiment of the invention provides a hydrocarbonconversion system comprising a plurality of reaction units, each unitcomprising a plurality of riser reactors, and at least one regenerationunit coupled to the reaction units. The number of reaction units isgreater than the number of regeneration units. Optionally, thisembodiment also provides at least one stripping unit coupled to thereaction units, wherein the number of reaction units is greater than thenumber of stripping units.

BRIEF DESCRIPTION OF THE DRAWINGS

[0017]FIG. 1 presents a partial cross sectional view of a hydrocarbonconversion apparatus of the present invention.

[0018]FIG. 2 presents a partial cross sectional view of anotherembodiment of the hydrocarbon conversion apparatus of the presentinvention.

[0019]FIG. 3 presents a partial cross sectional view of yet anotherembodiment of the hydrocarbon conversion apparatus of the presentinvention.

[0020]FIG. 4 presents a partial cross sectional view of still anotherembodiment of the hydrocarbon conversion apparatus of the presentinvention.

[0021]FIG. 5 presents cross sectional views of representativearrangements and configurations of the riser reactors and catalystreturns.

[0022]FIG. 6 presents a partial cross-sectional view of two multipleriser reactors and an integrated regeneration system in accordance withthe present invention.

[0023]FIG. 7 presents a partial cross-sectional view of two multipleriser reactors and an integrated regeneration system including anintegrated stripping system in accordance with the present invention.

DETAILED DESCRIPTION OF THE INVENTION

[0024] The present invention provides a method and a system forconverting a hydrocarbon over a catalyst to one or more products in amultiple reactor system while reducing the number of regeneration unitsand/or stripping units implemented therein. By reducing the number ofregeneration units and/or stripping units, the costs associated withmanufacturing and operating a multiple reactor system can be reduced.

[0025] During the catalytic conversion of hydrocarbons to variousproducts, e.g., the catalytic conversion of oxygenates to light olefins(the OTO process), carbonaceous deposits accumulate on the catalyst usedto promote the conversion reaction. At some point, the build up of thesecarbonaceous deposits causes a reduction in the capability of thecatalyst to function efficiently. For example, in the OTO process, anexcessively “coked” catalyst does not readily convert the oxygenate feedto light olefins. At this point, the catalyst is partially deactivated.When a catalyst can no longer convert the hydrocarbon to the desiredproduct, the catalyst is considered to be fully deactivated.

[0026] In accordance with the present invention, catalyst is withdrawnfrom a plurality of hydrocarbon conversion apparatuses (HCA's), e.g.,reactors or reaction units, and is directed to at least one regenerationunit. Preferably, at least one of the HCA's is a methanol to olefin(MTO) conversion apparatus. The catalyst is partially, if not fully,regenerated in the at least one regeneration apparatus. By regeneration,it is meant that the carbonaceous deposits are at least partiallyremoved from the catalyst. Desirably, the catalysts withdrawn from theHCA's are at least partially deactivated. The remaining portion of thecatalyst in the HCA's is re-circulated without regeneration, asdiscussed below. The regenerated catalyst, with or without cooling, isthen returned to the HCA's. Desirably, for each HCA, the rate ofwithdrawing the portion of the catalyst for regeneration is from about0.1% to about 99% of the rate of the catalyst exiting the reactor. Moredesirably, the rate is from about 0.2% to about 50%, and, mostdesirably, from about 0.5% to about 5%.

[0027] Optionally, the at least partially deactivated catalyst from theplurality of HCA's is directed to one or more stripping units whereinthe at least partially deactivated catalyst contacts one or morestripping mediums under conditions effective to recover adsorbedhydrocarbons from the at least partially deactivated catalyst.

[0028] As indicated above, the present invention is directed tocombining a plurality of HCA's with an integrated regeneration system.The plurality of HCA's could be selected from conventional HCA's and/ormultiple riser HCA's disclosed in more detail below with reference toFIGS. 1 through 5. The number of HCA's in fluid communication with theintegrated regeneration system depends on a variety of factors. Forexample, if a specific form of HCA is particularly effective in aspecific hydrocarbon conversion process, a plurality of those reactorsoptionally is coupled to a single regeneration system. In this manner,the amount of desired product produced can be increased because morethan one HCA is in use. Additionally, the cost of the overallhydrocarbon conversion process can be reduced because the number ofregeneration systems is reduced. The invention also provides for areduction in number of regeneration systems for a reactor system havingreactors that produce different products from one another, but which usethe same or very similar catalysts.

[0029] In one embodiment, the invention provides a reactor systemincluding a plurality of HCA's and a regenerator for converting an atleast partially deactivated catalyst to a regenerated catalyst. Thereactor system also includes a first conduit system for transferring theat least partially deactivated catalyst from the reactor units to theregenerator, and a second conduit system for transferring theregenerated catalyst from the regenerator to the plurality of reactorunits. The first conduit system optionally includes a first strippingunit for stripping the at least partially deactivated catalyst with afirst stripping medium. Also, the first conduit system optionallyincludes a second stripping unit for stripping the at least partiallydeactivated catalyst with a second stripping medium. The secondstripping medium can be the same as or different from the firststripping medium. Ideally, the first and second stripping units strip atleast partially deactivated catalysts from separate reactor units. Oneor more of the reactors preferably is a multiple-riser reactor.

[0030] In another embodiment, the invention provides a reactor systemcomprising a first reaction unit comprising a first plurality of riserreactors, and a second reaction unit comprising a second plurality ofriser reactors, wherein each of the first and second reaction units hasa first end into which a catalyst can be fed and a second end throughwhich the catalyst can exit the reaction unit. The reactor system alsoincludes a regeneration unit having a regeneration inlet and aregeneration outlet, and a regeneration line having a plurality of firstline ends in fluid communication with the second ends of the first andsecond reaction units and a second line end extending to theregeneration inlet. A return line is also provided having a first returnend in fluid communication with the regeneration outlet, a second returnend directing a first portion of the catalyst to the first reactionunit, and a third return end directing a second portion of the catalystto the second reaction unit.

[0031] Desirably, a portion of the catalyst, comprising molecular sieveand any other materials such as binders, fillers, etc., is removed fromeach HCA, e.g., reactor or reaction unit, for regeneration andrecirculation back to the HCA at a rate of from about 0.1 times to about10 times, more desirably from about 0.2 to about 5 times, and mostdesirably from about 0.3 to about 3 times the total feed rate ofoxygenates to the HCA. These rates pertain to the catalyst containingmolecular sieve only, and do not include non-reactive solids. The rateof total solids, i.e., catalyst and non-reactive solids, removed fromthe HCA for regeneration and recirculation back to the HCA will varythese rates in direct proportion to the content of non-reactive solidsin the total solids.

[0032] Desirably, the catalyst regeneration is carried out in one ormore regenerating units or regenerators in the presence of a gascomprising oxygen or other oxidants. Examples of other oxidants include,but are not necessarily limited to, singlet O₂, O₃, SO₃, N₂O, NO, NO₂,N₂O₅, and mixtures thereof. Air and air diluted with nitrogen or CO₂ aredesired regeneration gases. The oxygen concentration in air can bereduced to a controlled level to minimize overheating of, or creatinghot spots in, the regenerator. The catalyst can also be regeneratedreductively with hydrogen, mixtures of hydrogen and carbon monoxide, orother suitable reducing gases.

[0033] The catalyst can be regenerated in any number of methods—batch,continuous, semi-continuous, or a combination thereof. Continuouscatalyst regeneration is a desired method. Desirably, the catalyst isregenerated to a level of remaining coke from about 0.01 wt % to about15 wt %, more preferably from about 0.01 to about 5 wt %, of the weightof the catalyst.

[0034] The catalyst regeneration temperature should be from about 250°C. to about 750° C., and desirably from about 500° C. to about 700° C.Because the regeneration reaction preferably takes place at atemperature considerably higher, e.g., about 93° C. to about 150° C.higher, than the oxygenate conversion reaction, it is desirable to coolat least a portion of the regenerated catalyst to a lower temperaturebefore it is sent back to the reactor. One or more catalyst coolerslocated externally to the regeneration apparatuses can be used to removesome heat from the catalyst after it has been withdrawn from theregeneration apparatuses. When the regenerated catalyst is cooled, it isdesirable to cool it to a temperature which is from about 65° C. higherto about the temperature of the catalyst withdrawn from the reactor.This cooled catalyst is then returned to either some portion of thereactor, the regeneration apparatus, or both. When the regeneratedcatalyst from the regeneration apparatus is returned to the reactor, itcan be returned to any portion of the reactor. For example, the catalystcan be returned to the catalyst containment area to await contact withthe feed, the separation zone to contact products of the feed or acombination of both.

[0035] Desirably, catalyst regeneration is carried out after the atleast partially deactivated catalyst has been stripped of most of thereadily removable organic materials (organics), e.g., hydrocarbons, in astripper or stripping chamber. This stripping can be achieved by passinga stripping medium, e.g., a stripping gas over the spent catalyst at anelevated temperature. Gases suitable for stripping include steam,nitrogen, helium, argon, methane, CO₂, CO, hydrogen, and mixturesthereof. A preferred gas is steam. Gas hourly space velocity (GHSV,based on volume of gas to volume of catalyst and coke) of the strippinggas is from about 0.1 h⁻¹ to about 20,000 h⁻¹. Acceptable temperaturesof stripping are from about 250° C. to about 750° C., and desirably fromabout 400° C. to about 600° C.

[0036] An alternative embodiment of the invention is a method forforming olefins in a methanol to olefin (MTO) reactor system. The methodincludes contacting in a first reaction unit a first oxygenate with afirst catalyst under conditions effective to convert at least a portionof the first oxygenate to a first olefin and at least partiallydeactivating the first catalyst to form a deactivated first catalyst.The method also includes contacting in a second reaction unit a secondoxygenate with a second catalyst under conditions effective to convertat least a portion of the second oxygenate to a second olefin and atleast partially deactivating the second catalyst to form a deactivatedsecond catalyst. The deactivated first catalyst and deactivated secondcatalyst are directed to a regeneration unit and are regenerated to formregenerated catalysts. A first portion of the regenerated catalysts isdirected to the first reaction unit, and a second portion of theregenerated catalysts is directed to the second reaction unit. Thedeactivated first catalyst optionally contacts a first stripping mediumin a first stripping unit under conditions effective to removeinterstitial hydrocarbons from the deactivated first catalyst. Also, thedeactivated second catalyst optionally contacts a second strippingmedium in a second stripping unit under conditions effective to removeinterstitial hydrocarbons from the deactivated second catalyst.Alternatively, the deactivated second catalyst optionally contacts thefirst stripping medium in the first stripping unit under conditionseffective to remove interstitial hydrocarbons from the deactivatedsecond catalyst.

[0037] Another embodiment of the invention provides a hydrocarbonconversion system comprising first and second pluralities of riserreactors, each of the riser reactors having a first end into which acatalyst can be fed and a second end through which the catalyst can exitthe riser reactor. The hydrocarbon conversion system includes first andsecond catalyst retention zones provided to contain catalyst which canbe fed to the first and second plurality of riser reactors,respectively. Additionally, the system includes first and secondseparation zones into which the second ends of the first and secondpluralities of riser reactors extend, respectively, the separation zonesbeing provided to separate the catalyst from product(s) of a reaction orreactions conducted in the first and second pluralities of riserreactors. First and second catalyst returns are also provided in fluidcommunication between the first and second separation zones,respectively, and the first and second catalyst retention zones,respectively. The hydrocarbon conversion system also includes aregenerator for regenerating the catalyst, and first and second catalystoutlet lines, each of the outlet lines having a first end into which acatalyst can be fed from the first and second pluralities of riserreactors, respectively, and a second end through which the catalyst canenter the regenerator. The system also provides first and secondcatalyst return lines, each of the return lines having a first end intowhich a catalyst can be fed from the regenerator and a second endthrough which the catalyst can enter the first and second pluralities ofriser reactors, respectively.

[0038] Yet another embodiment of the invention provides a catalystregenerator system, comprising a regeneration zone for contacting an atleast partially deactivated catalyst with a regeneration medium underconditions effective to form a regenerated catalyst, a plurality ofcatalyst inlets for receiving the at least partially deactivatedcatalyst from a plurality of reactor units, and a plurality of catalystoutlets for delivering the regenerated catalyst to the plurality ofreactor units. A stripping zone or zones optionally is provided forcontacting the at least partially deactivated catalyst with a strippingmedium under conditions effective to remove interstitial hydrocarbonsfrom the deactivated catalyst.

[0039] In one embodiment, the invention is a method for regeneratingcatalyst comprising receiving an at least partially deactivated catalystfrom a plurality of multiple riser reaction units, heating the catalystunder conditions effective to convert the at least partially deactivatedcatalyst to a regenerated catalyst, and directing the regeneratedcatalyst to the plurality of multiple riser reaction units.

[0040] Another embodiment of the invention provides a hydrocarbonconversion system comprising a plurality of reaction units, each unitcomprising a plurality of riser reactors, and at least one regenerationunit coupled to the reaction units. The number of reaction units isgreater than the number of regeneration units. Optionally, thisembodiment also provides at least one stripping unit coupled to thereaction units, wherein the number of reaction units is greater than thenumber of stripping units. Optionally, this embodiment also provides atleast one catalyst cooler coupled to the reaction units, wherein thenumber of reaction units is greater than the number of catalyst coolers.

[0041] In another embodiment of the present invention, a plurality ofHCA's for converting hydrocarbons to different products, e.g., one ormore of the HCA's producing different products from the other HCA's, iscoupled to a single regeneration system. This embodiment is particularlyeffective where the catalysts implemented in the various hydrocarbonconversion processes are the same or similar for the differenthydrocarbon conversion processes. In one particularly preferredembodiment a plurality of the HCA's are MTO conversion apparatuses.Optionally, a first MTO conversion apparatus can be designed to providea first ethylene-to-propylene product ratio. The first ratio may bedifferent from a second ethylene-to-propylene ratio provided by a secondMTO conversion apparatus. Thus, a first MTO conversion apparatus mayproduce mostly ethylene while a second MTO conversion apparatus mayproduce mostly propylene, although the two MTO conversion apparatusesshare a common regeneration system.

[0042] As indicated above, the number of HCA's that are coupled to asingle regeneration system varies. For example, the present inventionprovides for two, three, four, five, six, seven, eight, nine, ten ormore HCA's in fluid communication with a lesser number of regenerationsystems. Preferably, a plurality of HCA's are in fluid communicationwith a single regeneration system.

[0043] As used herein, “regeneration system” means one or moreregeneration units, one or more stripping units, and/or one or morecatalyst coolers for cooling the catalyst prior to recycling theregenerated catalyst to the HCA's. The “regeneration system” optionallyincludes various conduits or lines coupling these units, e.g., theHCA's, the stripper(s), the regeneration unit(s), and/or the catalystcooler(s).

[0044]FIG. 6 illustrates one embodiment of an integrated regenerationsystem, generally designated 600, in accordance with one embodiment ofthe present invention. The integrated regeneration system 600 is coupledto a plurality of HCA's 602A, 602B. Although FIG. 6 only illustrates twoHCA's in fluid communication with the integrated regeneration system600, in other embodiments, the integrated regeneration system is also influid communication with additional HCA's, as discussed above. As shown,the HCA's 602A, 602B are substantially similar or identical in form.Alternatively, the HCA's are different from each other. As shown in FIG.6, in one embodiment, each HCA 602A, 602B includes multiple riserreactors. The feedstock 630A, 630B is shown entering the bottom of eachof the HCA's 602A, 602B. In the HCA's, the feedstock contacts catalystunder conditions effective to convert at least a portion of thefeedstock to product. An HCA product effluent stream 634 is shownexiting the separation zones of each of the HCA's. The product effluentlines from each HCA are combined and then directed to a product recoveryunit (not shown).

[0045] The process of the integrated regeneration system 600 will now bedescribed in more detail. The separation zones of the HCA's preferablyat least partially separate catalyst from the desired product. At leasta portion of the catalyst is then withdrawn from each HCA. The withdrawncatalyst can include partially deactivated, fully deactivated and/oractivated catalysts, e.g., containing substantially no carbonaceousdeposits.

[0046] With continuing reference to FIG. 6, at least a portion of thecatalyst from the HCA's 602A, 602B is withdrawn through conduits orlines 604A, 604B. The lines 604A, 604B optionally include one or moreflow control devices 606A, 606B. Each flow control device 606A, 606B canbe of any type of flow control device currently in use in the art tocontrol catalyst flow through the catalyst transfer lines. Usefulnon-limiting examples of flow control devices include ball valves, plugvalves and slide valves. Preferably, the present invention includes oneor more strippers or stripping units 608A, 608B. As shown in FIG. 6,each HCA has its own respective stripping unit 608A, 608B.

[0047] In this embodiment, the catalyst from the plurality of HCA's602A, 602B is directed to the stripping units 608A, 608B. In thestripping unit, the at least partially deactivated catalyst contacts astripping medium, which enters the stripping units through lines 632A,632B, under conditions effective to remove at least a portion of theorganics, e.g., interstitial hydrocarbons, from the catalyst. Thestripped catalyst then is withdrawn through lines 610A and 610B andoptionally passes through one or more flow control devices 612A, 612B.Stripper offgas streams 636A, 636B, comprising the stripping medium,light hydrocarbons and desirable product, exit the top of the strippingunits 608A, 608B and optionally are directed to and combined withproduct effluent stream 634. In an alternative embodiment (not shown),stripper offgas streams 636A, 636B are directed to the HCA's 602A, 602B,e.g., to the separation zones thereof, wherein entrained catalyst can beseparated from the product in the offgas streams.

[0048] The stripped catalyst from lines 610A, 610B is then directed to asingle regeneration unit 614. Optionally, lines 610A and 610B fromstripping units 608A and 608B, respectively, can be combined into asingle line (not shown) for transferring the combined catalyst streamfrom HCA's 602A and 602B into regeneration unit 614. Because theregeneration unit 614 receives catalyst from a plurality of HCA's, thesize of the regeneration unit preferably is adapted to be larger thanconventional regeneration units in order to accommodate the largervolume of catalyst being received therein. A regenerating medium, whichpreferably has been compressed, such as air is provided to theregeneration unit 614 through lines 628. The regenerating mediumcontacts the stripped catalyst in the regeneration unit under conditionseffective to at least partially regenerate the stripped catalyst.Regenerator offgas line 638 removes byproducts of the regenerationprocess, e.g., N₂, CO₂, and CO, from the regeneration unit 614. Afterthe catalyst has been regenerated, it is withdrawn from regenerationunit 614 via catalyst outlet line 616 and optionally passes through aflow control device 618. At dividing point 620, the regenerated catalyststream is divided between a plurality of lines adapted to carry theregenerated catalyst to one or more catalyst coolers 640A, 640B (two areillustrated), wherein the regenerated catalyst contacts a coolingmedium, e.g., water or steam, under conditions effective to cool theregenerated catalyst. In an alternative embodiment, not shown, theregeneration unit 614 can include a plurality of catalyst outlet lines616, each of which is directed to a respective catalyst cooler. Each ofthe plurality of lines that directs the regenerated catalyst to therespective catalyst coolers 640A, 640B preferably includes a flowcontrol device 624A, 624B. A first portion of the cooled catalyst fromthe catalyst coolers 640A, 640B is redirected through lines 642A, 642Bback to the regeneration unit 614 for further regeneration. A secondportion of the cooled catalyst is directed through lines 622A, 622B tocatalyst inlets 626A, 626B of respective HCA's 602A, 602B. Preferably,each of lines 622A and 622B includes a flow control device 644A, 644Bfor controlling the flow of catalysts therein.

[0049]FIG. 7 illustrates another embodiment of the present inventionwherein catalyst from a plurality of HCA's is directed to a singlestripping unit 702 and a single catalyst cooler 716. Specifically, atleast partially deactivated catalyst is withdrawn from the HCA's 602Aand 602B via lines 704A and 704B. Each of these lines optionallyincludes a flow control device 706A, 706B. Lines 704A and 704B carry theat least partially deactivated catalyst from the HCA's to the singlestripping unit 702. As shown in FIG. 7, the at least partiallydeactivated catalysts from the HCA's are directed separately from theHCA's to the stripping unit 702. However, in another embodiment (notillustrated), lines 704A and 704B are combined into a single feed linewhich introduces the combined catalyst stream from the HCA's intostripping unit 702. Once in the stripping unit 702, the catalystcontacts a stripping medium, which enters the stripping unit throughline 712, under conditions effective to remove interstitial hydrocarbonstherefrom. Stripper offgas stream 714 comprising the stripping medium,light hydrocarbons and desirable product, exits the top of the strippingunit 702 and optionally is directed to and combined with producteffluent stream 634. In an alternative embodiment (not shown), stripperoffgas stream 714 is directed to the HCA's 602A, 602B, e.g., to theseparation zones thereof, wherein entrained catalyst can be separatedfrom the product in the offgas streams.

[0050] With further reference to FIG. 7, after the catalyst has beenstripped in unit 702, it is withdrawn from stripping unit 702 via line708 and optionally passes through a flow control device 710 in route tothe regeneration unit 614. In the regeneration unit, the strippedcatalyst contacts a regenerating medium, which has been introduced intothe regeneration unit 614 through lines 628, under conditions effectiveto regenerate at least a portion of the catalysts. After regeneration inregeneration unit 614, the regenerated catalyst is withdrawn throughline 616, passes through a flow control device 618, and is directed to acatalyst cooler 716. In the catalyst cooler, the catalyst contacts acooling medium under conditions effective to cool the catalyst. A firstportion of the cooled catalyst is redirected through line 718 back tothe regeneration unit 614 for further regeneration. Second and thirdportions of the cooled catalyst are directed, respectively, throughlines 622A and 622B, flow control devices 624A and 624B, and back to theHCA's 602A, 602B via inlets 626A, 626B.

[0051] The embodiment disclosed in FIG. 7 provides the additional costsaving advantage of stripping catalyst from a plurality of HCA's in asingle stripping unit. Thus, as with regeneration unit 614, the size ofstripping unit 702 optionally is larger than a conventional strippingunit associated with a single reactor system in order to accommodate theincreased volume of catalyst received therein. The embodiment disclosedin FIG. 7 also provides the cost saving advantage of cooling catalystfrom a regeneration unit a single catalyst cooler. Thus, as withregeneration unit 614, the size of catalyst cooler 716 optionally islarger than a conventional catalyst cooler associated with a singlereactor system in order to accommodate the increased volume of catalystreceived therein. In another embodiment, a plurality of catalyst coolersis implemented in accordance with the present invention for cooling eachrespective stream of regenerated catalysts prior to its introduction toeach respective HCA.

[0052] As shown in FIGS. 6 and 7, the catalyst is withdrawn fromregeneration unit 614 via a single line 616. However, in an alternativeembodiment (not shown), the regenerated catalyst is withdrawn from theregeneration unit via a plurality of lines, each of which carries theregenerated catalyst to a respective catalyst cooler and/or HCA.Nevertheless, withdrawing regenerated catalysts from a single line ispreferred in accordance with the present invention so that a singlecatalyst cooler can cool the regenerated catalyst prior to itsreintroduction into the plurality of HCA's, as shown in FIG. 7.

[0053] Additionally, although the present invention has generally beendescribed as directing catalyst from a plurality of HCA's to a singlestripping unit and/or a single regeneration unit and/or a singlecatalyst cooler, benefits can also be realized from directing catalystfrom a plurality of HCA's to a plurality of stripping units,regeneration units and/or catalyst coolers. For example, in analternative embodiment of the present invention, catalysts from aplurality of HCA's are directed to at least one regeneration unitcoupled thereto, wherein the number of HCA's, e.g., reaction units, isgreater than the number of regeneration units. Similarly, catalyst froma plurality of HCA's is directed to at least one stripping unit, whereinthe number of HCA's is greater than the number of stripping units.Additionally or alternatively, regenerated catalysts is directed to atleast one catalyst cooler wherein the number of HCA's is greater thanthe number of catalyst coolers.

[0054] The process for converting a feedstock, especially a feedstockcontaining one or more oxygenates, in the presence of a molecular sievecatalyst composition of the invention, is carried out by a reactionprocess in a plurality of HCA's where the processes are preferablyselected from the following types of processes: a fixed bed process, afluidized bed process (including a turbulent bed process), preferably acontinuous fluidized bed process, and most preferably a continuous highvelocity fluidized bed process. In accordance with the presentinvention, one or more reaction processes, that is, the same ordifferent processes, occurs in a plurality of HCA's.

[0055] The types of individual HCA's that can be implemented in thepresent invention will now be described in more detail. The reactionprocesses can take place in a variety of catalytic reactors such ashybrid reactors that have dense bed or fixed bed reaction zones and/orfast fluidized bed reaction zones coupled together, circulatingfluidized bed reactors, riser reactors, and the like. Suitableconventional reactor types are described in for example U.S. Pat. No.4,076,796, U.S. Pat. No. 6,287,522 (dual riser), and FluidizationEngineering, D. Kunii and 0. Levenspiel, Robert E. Krieger PublishingCompany, New York, N.Y. 1977, which are all herein fully incorporated byreference.

[0056] The preferred reactor type is selected from the riser reactorsgenerally described in Riser Reactor, Fluidization and Fluid-ParticleSystems, pages 48 to 59, F. A. Zenz and D. F. Othmer, ReinholdPublishing Corporation, New York, 1960, and U.S. Pat. No. 6,166,282(fast-fluidized bed reactor), which are incorporated herein byreference. Most preferably, at least one HCA includes a plurality ofriser reactors, as disclosed in U.S. patent application Ser. No.09/564,613 filed May 4, 2000 to Lattner et al., the entirety of which isincorporated herein by reference, and as discussed in more detail below.Optionally, each of the HCA's includes a plurality of riser reactors, asshown in FIGS. 6 and 7.

[0057] In the preferred embodiment, a fluidized bed process or highvelocity fluidized bed process includes a reactor system, a regenerationsystem and a recovery system. The reactor system preferably includes aplurality of fluid bed reactor systems, each having a first reactionzone within one or more riser reactor(s) and a second reaction zonewithin at least one disengaging vessel, preferably comprising one ormore cyclones. In one embodiment, the one or more riser reactor(s) and adisengaging vessel is contained within a single HCA. Fresh feedstock,preferably containing one or more oxygenates, optionally with one ormore diluent(s), is fed to the one or more riser reactor(s) in a givenHCA in which a zeolitic or non-zeolitic molecular sieve catalystcomposition or coked version thereof is introduced. In one embodiment,the molecular sieve catalyst composition or coked version thereof iscontacted with a liquid or gas, or combination thereof, prior to beingintroduced to the riser reactor(s), preferably the liquid is water ormethanol, and the gas is an inert gas such as nitrogen.

[0058] In one embodiment, the amount of liquid feedstock fed separatelyor jointly with a vapor feedstock, to a reactor system is in the rangeof from 0.1 weight percent to about 95 weight percent, preferably fromabout 10 weight percent to about 90 weight percent, more preferably fromabout 50 weight percent to about 85 weight percent based on the totalweight of the feedstock including oxygenate recycle and any diluentcontained therein. The liquid and vapor feedstocks are preferably thesame composition, or contain varying proportions of the same ordifferent feedstock with the same or different diluent.

[0059] The feedstock entering an individual reactor system is preferablyconverted, partially or fully, in the first reactor zone into a gaseouseffluent that enters the disengaging vessel along with a coked molecularsieve catalyst composition which is at least partially deactivated. Inthe preferred embodiment, cyclone(s) within the disengaging vessel aredesigned to separate the molecular sieve catalyst composition,preferably a coked molecular sieve catalyst composition, from thegaseous effluent containing one or more olefin(s) within the disengagingzone. Cyclones are preferred, however, gravity effects within thedisengaging vessel will also separate the catalyst compositions from thegaseous effluent. Other methods for separating the catalyst compositionsfrom the gaseous effluent include the use of plates, caps, elbows, andthe like.

[0060] In one embodiment of a disengaging system in an individual HCA,the disengaging system includes a disengaging vessel. In one embodiment,a lower portion of the disengaging vessel is a stripping zone. In thestripping zone the at least partially coked molecular sieve catalystcomposition is contacted with a stripping medium which is a gas,preferably one or a combination of steam, methane, carbon dioxide,carbon monoxide, hydrogen, or an inert gas such as argon, preferablysteam, to recover adsorbed interstitial hydrocarbons from the at leastpartially coked molecular sieve catalyst composition that is thenintroduced to the regeneration system. Ideally, from about 2 to about10, more preferably about 2 to about 6, and most preferably 3 to about 5pounds of stripping medium, e.g., steam, is provided to the strippingunit for every 1000 pounds of catalyst. In another embodiment, thestripping zone is in a separate vessel from the disengaging vessel andthe stripping medium is passed at a gas hourly superficial velocity(GHSV) of from 1 hr⁻¹ to about 20,000 hr⁻¹ based on the volume of gas tovolume of coked molecular sieve catalyst composition, preferably at anelevated temperature from 250° C. to about 750° C., preferably fromabout 350° C. to 650° C., over the coked molecular sieve catalystcomposition.

[0061] In one embodiment, catalyst from more than one disengagingsystem, e.g., from more than one HCA, is directed to an integratedstripping zone.

[0062] In one preferred embodiment of the process for converting anoxygenate to olefin(s) using a silicoaluminophosphate molecular sievecatalyst composition, the process is operated at a WHSV of at least 20hr⁻¹ and a Temperature Corrected Normalized Methane Selectivity (TCNMS)of less than 0.016, preferably less than or equal to 0.01. See, forexample, U.S. Pat. No. 5,952,538, which is herein fully incorporated byreference.

[0063]FIG. 1 presents a partial cross sectional view of a HCA 10 thatoptionally is implemented with the integrated regeneration system inaccordance with the present invention. The apparatus 10 comprises ashell 12, a plurality of riser reactors 20, a feed distributor 30, and acatalyst return 50. Preferably, the present invention couples aplurality of HCA's to an integrated regeneration system.

[0064] With continuing reference to FIG. 1, the shell 12 forms aseparation zone 14 in which a product of the hydrocarbon conversionreaction is separated from the catalyst which catalyzes the hydrocarbonconversion reaction. Shell 12 includes a first end 16 and a second end18. The separation zone 14 may additionally contain one or moreseparation devices, not shown, which are used to separate the productsfrom the catalyst. Useful separation devices are discussed below inassociation with the discussion of other embodiments of the presentinvention. Further, the separation devices may be positioned externallyto the separation zone 14, i.e., outside of the shell 12 of the HCA 10,or a combination of externally and internally positioned separationdevices.

[0065] Optionally, the riser reactors 20 in one or more of the HCA's mayextend into shell 12 and into the separation zone 14. By extending theriser reactors 20 into shell 12 and the separation zone 14, the heightrequired to obtain the desired aspect ratio of a given riser reactor 20is concurrent with at least a portion of the height required for theshell 12, separation zone 14, and other associated spaces, reducing thetotal height of the hydrocarbon conversion reactor 10 of the presentinvention. Each riser reactor 20 includes a first end 22 into which thecatalyst and feed are fed to conduct the hydrocarbon conversionreaction. Each riser reactor 20 further includes a second end 24 throughwhich the catalyst, products and unreacted feed, if any, exit the riserreactor 20. The first end 22 of each riser reactor 20 terminates in amouth 26 through which the catalyst and feed are fed into the riserreactor 20. The number of riser reactors 20 employed in each HCA 10varies depending on the hydrocarbon conversion process or processes tobe conducted in the plurality of apparatuses 10. Each apparatus 10 cancontain two, three, four, five, six or even more than six riser reactors20.

[0066] In another embodiment (not shown), one or more of the HCA'sinclude a transport conduit for consolidating and directing the outputof multiple reactors to the separation zone, as disclosed in U.S. patentapplication Ser. No. 09/564,613 to Lattner et al. (multiple riserreactor), the entirety of which is incorporated herein by reference.

[0067] The size of the riser reactors 20 depends on parameters such asthe superficial gas velocity, solids hydrodynamics, pressure, andproduction capacity of the desired hydrocarbon conversion process. Inthe present invention, each riser reactor 20 desirably has a height from10 meters to 70 meters and a width (or diameter) of one meter to threemeters. All of the riser reactors 20 in a given HCA have a similarheight from their first ends 22 to their second ends 24. Desirably, theheights of the riser reactors 20 vary by no more than 20% from one riserreactor 20 to another riser reactor 20 in an individual HCA. Moredesirably, the heights vary by no more than 10% and, most desirably, theheights vary by no more than 1%.

[0068] In the present invention, each of the riser reactors 20 in anindividual HCA has a similar cross sectional area along its entireheight. Desirably, each of the riser reactors 20 has a cross sectionalarea of no greater than 12 m². More desirably, each of the riserreactors 20 has a cross sectional area of no greater than 7 m². Mostdesirably, each of the riser reactors 20 has a cross sectional area ofno greater than 3.5 m². Desirably, the cross sectional areas of theriser reactors 20 vary by no more than 20% from one riser reactor 20 toanother riser reactor 20. More desirably, the cross sectional areas ofthe riser reactors 20 vary by no more than 10% and, most desirably, thecross sectional areas of the riser reactors 20 vary by no more than 1%.If one or more riser reactors 20 have both a largest and a smallestcross-sectional area at different points along the height of riserreactors 20, desirably the largest cross-sectional areas of the riserreactors 20 vary by no more than 20% from one riser reactor 20 toanother riser reactor 20, and the smallest cross-sectional areas of theriser reactors 20 vary by no more than 20% from one riser reactor 20 toanother riser reactor 20. More desirably, the largest cross sectionalarea of one riser reactor 20 varies by no more than 10% from the largestcross sectional area of another riser reactor 20 and the smallest crosssectional area varies by no more than 10% from the smallest crosssectional area of another riser reactor 20. Most desirably, the largestcross sectional area of one riser reactor 20 varies by no more than 1%from the largest cross sectional area of another riser reactor 20 andthe smallest cross sectional area varies by no more than 1% from thesmallest cross sectional area of another riser reactor 20.

[0069] Desirably, the cross sectional area of each riser reactor 20 inan individual HCA varies by no more than 50% along its entire length.More desirably, the cross sectional area of each riser reactor 20 in anHCA varies by no more than 30% along its entire height and, mostdesirably, the cross sectional area of each riser reactor 20 varies byno more than 10% along its entire height.

[0070] To provide a feed to the riser reactors 20 of an HCA, at leastone feed distributor 30 is positioned near the first ends 22 of theriser reactors 20. More than one feed distributor 30 may be employedadjacent the first ends 22 of the riser reactors 20 to provide feed invarious states, e.g., one feed distributor 30 may provide feed in avapor form while a second feed distributor 30 may provide feed in aliquid form. Feed distributor 30 includes a body 32 from which aplurality of necks 34 extend. Each riser reactor 20 has at least oneassociated neck 34. Each neck 34 terminates in a head 36. Each head 36of each neck 34 is positioned adjacent to the first end 22 of each riserreactor 20. Desirably, each head 36 extends upwardly into each riserreactor 20. More desirably, each head 36 is positioned at or above themouth 26 at the first end 22 of each riser reactor 20. Feed distributor30 may include an optional flow control device, not shown, positioned onfeed distributor 30 to control the amount of feed to each neck 34 or aflow control device may be positioned on each neck 34. The flow controldevice can also be employed to measure flow as well as control it.Further, a nozzle, not shown, may be positioned on each head 36 tofurther control the distribution of the feed to each riser reactor 20.Additionally, each head 36 may be fitted with screening device, notshown, to prevent back flow of catalyst into any of necks 34 of feeddistributor 30 and, subsequently, into body 32 of feed distributor 30.

[0071] At least one catalyst return 50 provides fluid communicationbetween the separation zone of 14 of shell 12 and the riser reactors 20.Particularly, each catalyst return 50 provides fluid communicationbetween the separation zone 14 and the first ends 22 of each riserreactor 20. Each catalyst return 50 has a first end 52 and a second end54. The first end 52 of the catalyst return 50 opens into the second end18 of shell 12 and the second end 54 of catalyst return 50 opensadjacent the riser reactors 20. Each catalyst return 50 is provided totransport catalyst from the separation zone 14 of shell 12 to the firstends 22 of the riser reactors 20. One or more of the HCA's may includeone, two, three, four, five, six or more catalyst returns 50. Typically,although not necessarily, the number of catalyst returns 50 in an HCAcorresponds to the number of riser reactors 20 in an individual HCA. Inthe embodiment shown in FIG. 1, the catalyst returns 50 are external tothe riser reactors 20. However, as shown in subsequently describedembodiments, the catalyst return 50 may be contained within a commonshell or be positioned internally in relation to the riser reactors 20or some combination thereof. Flow of catalyst through the catalystreturn(s) 50 may optionally be controlled through the use of a flowcontrol device 56 positioned on each catalyst return 50. The flowcontrol device 56 can be any type of flow control device currently inuse in the art to control catalyst flow through catalyst transfer lines.If employed, the flow control device 56 is desirably a ball valve, aplug valve or a slide valve.

[0072] The HCA 10 further includes a base 60. In the embodiment shown inFIG. 1, the base 60, the catalyst returns 50 and the first ends 22 ofthe riser reactors 20 define a catalyst retention zone 62. The catalystretention zone 62 is provided to retain catalyst which is used tocatalyze the hydrocarbon conversion reaction which is conducted in theapparatus 10. The catalyst return 50 provides fluid communicationbetween the separation zone 14 and the catalyst retention zone 62. To doso, the second ends 54 of the catalyst returns 50 open to the catalystretention zone 62. As one of skill in the art will appreciate, theboundary between the catalyst retention zone 62 and the catalyst return50 is fluid and depends, at least in part, on the level of catalystcontained in the catalyst return 50 and the catalyst retention zone 62.

[0073] A fluid distributor 70 is also positioned in or near the base 60of the apparatus 10. The fluid distributor 70 includes a conduit 72 intowhich a fluidizing fluid is fed into catalyst retention zone 62 tofluidize a fluidizable catalyst in the catalyst retention zone 62 andthe catalyst returns 50. Additional fluid distributors 70, as shown inFIG. 1, may also be positioned on each catalyst return 50 to fluidize afluidizable catalyst contained in each of the catalyst returns 50.

[0074] One or more of the HCA's may also include an outlet 80 throughwhich the catalyst can be removed from the apparatus 10. The outlet 80is shown as being positioned on the second end 18 of the shell 12 butmay be positioned at any position on the apparatus 10. The apparatus 10may also include an inlet 82 through which the catalyst may be placedinto the apparatus 10. Although the inlet 82 is shown as beingpositioned on the first end 16 of the shell 12, the inlet 82 may bepositioned at any position on the apparatus 10. A line 84 may beprovided to remove hydrocarbon conversion products from the apparatus10.

[0075] As shown in FIG. 1, the present invention preferably includes anassociated catalyst regeneration apparatus 90. The catalyst regenerationapparatus 90 is in fluid communication with the HCA 10. The catalystregeneration apparatus 90 includes a catalyst regenerator 92, which isin fluid communication with the HCA 10, and, optionally catalyststripper 94, which is in fluid communication with the catalystregenerator 92 and which may be in fluid communication with one or moreof the HCA's. A first line 96 provides fluid communication between theoutlet 80 on shell 12 and the catalyst stripper 94. A second line 98provides fluid communication between the catalyst stripper 94 and thecatalyst regenerator 92. A third line 100 provides fluid communicationbetween the catalyst regenerator 92 and the inlet 82 on shell 12. A flowcontrol device 102 may optionally be positioned on first line 96 tocontrol the flow of catalyst between the shell 12 and the catalyststripper 94. A flow control device 104 may optionally be positioned onsecond line 98 to control the flow of catalyst between the catalyststripper 94 and the catalyst regenerator 92. Although the catalyststripper 94 is shown on FIG. 1 as being separate from the catalystregenerator 92, one skilled in the art will appreciate that the catalyststripper 94 may be integrally formed with the catalyst regenerator 92.One skilled in the art will also appreciate that, although FIG. 1 showsthird line 100 as returning the catalyst to the separation zone 14through line 82, the catalyst may also be returned to the catalystreturn 50, the catalyst retention zone 62 and combinations of theseparation zone 14, the catalyst return 50 and the catalyst retentionzone 62.

[0076] When in operation, one or more of the HCA's, as shown in FIG. 1,function in the following manner. The apparatus 10 is filled with anappropriate amount of a catalyst suitable to carry out the desiredhydrocarbon conversion reaction. The catalyst should be of a type whichis fluidizable. At least a portion of the catalyst is contained in thecatalyst retention zone 62. To fluidize the catalyst in the catalystretention zone 62, a fluidizing fluid is fed into the fluiddistributor(s) 70 through inlet 72. The fluidizing fluid is fed into thecatalyst retention zone 62 and the catalyst return(s) 50 of the HCA 10.Useful fluidizing fluids include, but are not limited to, inert gasses,nitrogen, steam, carbon dioxide, and hydrocarbons. The choice offluidizing fluid depends upon the type of conversion reaction beingconducted in the HCA 10. Desirably, the fluidizing fluid is unreactive(i.e. inert) in the reaction being conducted in the HCA 10. In otherwords, it is desirable that the fluidizing fluid does not play a part inthe hydrocarbon conversion process being conducted in the HCA 10 otherthan to fluidize the fluidizable catalyst.

[0077] Once the catalyst has reached an acceptable fluidized state, afeed is fed into the HCA 10 through feed distributor 30. The feed entersthe body 32 of feed distributor 30, passes through the necks 34 of feeddistributor 30 and exits through the heads 36 of feed distributor 30.The feed is distributed to each of the riser reactors 20 through theirfirst ends 22. Desirably, the feed is provided in substantially equalstreams to each riser reactor 20. By “substantially equal” it is meantthat the flow of feed provided to each riser reactor 20 through the feeddistributor 30 varies by no more than 25% by volume rate, and varies nomore than 25% by mass percent for each component in the feed, from oneriser reactor 20 to another riser reactor 20. More desirably, the flowof feed provided to each riser reactor 20 through the feed distributor30 varies by no more than 10% by volume rate, and varies no more than10% by mass percent for each component in the feed, from one riserreactor 20 to another riser reactor 20. Most desirably, feed provided toeach riser reactor 20 through the feed distributor 30 varies by no morethan 1% by volume rate, and varies no more than 1% by mass percent foreach component in the feed, from one riser reactor 20 to another riserreactor 20.

[0078] A pressure differential created by the velocity of the feedentering the first ends 22 of the riser reactors 20 and the pressure ofthe height of fluidizable catalyst in the catalyst return(s) 50 and thecatalyst retention zone 62 causes catalyst to be aspirated into thefirst ends 22 of the riser reactors 20. The catalyst is transportedthrough the riser reactors 20 under well known principles of eduction inwhich the kinetic energy of one fluid, in this case the feed, is used tomove another fluid, in this case the fluidized catalyst. The catalystand feed travel from the first ends 22 to the second ends 24 of theriser reactors 20. As the catalyst and feed travel through the riserreactors 20, the hydrocarbon conversion reaction occurs and a conversionproduct is produced.

[0079] By designing one or more HCA's with these features, eachindividual riser reactor 20 in a given HCA operates in a substantiallyidentical manner. With this invention, it is desirable to maintain boththe reactant feed rates and the catalyst feed rates at the same rates toeach of the riser reactors 20. In this way, the conversion of the feedand selectivity to the desired products will be substantially identicaland can run at optimum operational conditions.

[0080] The conversion product(s), unreacted feed, if any, and thecatalyst exit the riser reactors 20 through their second ends 24 andenter into the separation zone 14 of shell 12. In second end 16 of shell12, the conversion product and unreacted feed, if any, are separatedfrom the catalyst by a separator, not shown, such as cyclonicseparators, filters, screens, impingement devices, plates, cones, otherdevices which would separate the catalyst from the product of theconversion reaction, and combinations thereof. Desirably, the conversionproduct and unreacted feed, if any, are separated by a series ofcyclonic separators. Once the catalyst has been separated from theconversion product and the unreacted feed, if any, the conversionproducts and unreacted feed, if any, are removed from the shell 12through the line 84 for further processing such as separation andpurification. The catalyst, after being separated from the products andunreacted feed, moves from the shell 12 to the catalyst retention zone62. The catalyst exits shell 12 through the first ends 52 of thecatalyst returns 50 and moves through the catalyst returns 50 to thefirst ends 54 of the catalyst returns 50 from which the catalyst movesto the catalyst retention zone 62. If desired, the flow of catalystthrough the catalyst returns 50 can be controlled by the flow controldevices 56. If the flow control devices 56 are employed, a height offluidizable catalyst is maintained above each flow control device 56 inthe catalyst return 50 to allow proper function of the flow controldevice 56.

[0081] In accordance with the present invention, at least a portion ofthe catalyst from a plurality of HCA's is circulated to a catalystregeneration apparatus 90, as shown in FIG. 1. Catalyst to beregenerated is removed from the shell 12 though the outlet 80 andtransported, if desired, to the catalyst stripper 94 through the firstline 96. Optionally, the catalyst stripper 94 may include a second inputline (not shown) from a second HCA (not shown). In this manner, catalystfrom the second HCA can be transported by the second input line to thecatalyst stripper 94. The flow of catalyst from the HCA 10 to thecatalyst stripper 94 can be controlled by the flow control device 102.In the catalyst stripper 94, the catalyst is stripped of most of readilyremovable organic materials (organics), e.g., hydrocarbons. Strippingprocedures and conditions for individual hydrocarbon conversionprocesses are within the skill of a person of skill in the art. Thestripped catalyst is transferred from the catalyst stripper 94 to thecatalyst regenerator or regeneration unit 92 through the second line 98.The flow of catalyst through the second line 98 may optionally becontrolled by the optional flow control device 104. In the catalystregenerator 92, carbonaceous deposits (coke) formed on the catalystduring a hydrocarbon conversion reaction are at least partially removedfrom the catalyst.

[0082] The catalyst regenerator 92 preferably receives catalyst from aplurality of HCA's as illustrated in FIG. 6, and as discussed in moredetail below. As shown in FIG. 1, the catalyst regenerator 92 optionallymay receive catalyst via second line 98 and via line 101, which isconnected to a second HCA (not shown). Optionally, line 101 may bedirected to a second stripping unit (not shown) which is in fluidcommunication with the second HCA. In this embodiment, catalyst may bedelivered to the catalyst regenerator 92 from the second HCA throughline 101. Similarly, catalyst from more than two HCA's may be directedto the catalyst stripper 94. Additionally or alternatively, catalystfrom more than two HCA's may be directed to the catalyst regenerator 92.

[0083] The regenerated catalyst is then transferred to the shell 12 ofthe HCA 10 through outlet line 111 and third line 100. Regeneratedcatalyst may also be directed to other hydrocarbon apparatuses coupledto the regeneration system. A transport gas is typically provided to theoutlet line 111, the third line 100, and line 105 to facilitate transferof the catalyst from the catalyst regenerator 92 to the HCA 10 and thesecond HCA (not shown) and any additional HCA's (HCA's). As shown inFIG. 1, the catalyst is returned to the shell 12 through the inlet 82.

[0084] The catalyst regenerator 92 in accordance with the presentinvention also provides a conduit system for returning at leastpartially regenerated catalyst to the second hydrocarbon conversionreactor (not shown). This objective may be realized as shown in FIG. 1whereby regenerated catalyst exits the catalyst regenerator in lineoutlet line 111 and is divided between two lines, third line 100 andline 105, at dividing point 109. The flow of catalyst through the outletline 111, the third line 100, and line 105 may optionally be controlledby flow control devices 106, 103 and 107, respectively. The flow controldevices 102, 103, 104, 106, and 107 can be any types of flow controldevices currently in use in the art to control catalyst flow throughcatalyst transfer lines. Useful non-limiting flow control devicesinclude ball valves, plug valves and slide valves.

[0085] One or more of the HCA's may be adapted as shown in FIG. 2. Inthis alternative embodiment, HCA 110, shown in partial cross section,comprises a shell 120, a plurality of riser reactors 130, a feeddistributor 140, and a catalyst return 150.

[0086] With continuing reference to FIG. 2, the shell 120 forms aseparation zone 122 in which a product of the hydrocarbon conversionreaction is separated from the catalyst which catalyzes the hydrocarbonconversion reaction. Shell 120 includes a first end 124 and a second end126. Shell 120 defines a quiescent zone 128 from which catalyst can bewithdrawn from the HCA 110.

[0087] Riser reactors 130 in a given HCA extend into shell 120 and theseparation zone 122. Each riser reactor 130 includes a first end 132into which the catalyst and feed are fed to conduct the hydrocarbonconversion reaction. Each riser reactor 130 further includes a secondend 134 through which the catalyst, products and unreacted feed, if any,exit the riser reactor 130. The first end 132 of each riser reactor 130terminates in a mouth 136 through which the catalyst and feed are fedinto the riser reactor 130. As described above, the number of riserreactors 130 employed in the HCA 110 varies depending on the hydrocarbonconversion process to be conducted in the apparatus 110. The number andsize of the riser reactors 130 is discussed above in conjunction withthe description of FIG. 1.

[0088] To provide a feed to the riser reactors 130, at least one feeddistributor 140 is positioned near the first ends 132 of the riserreactors 130. More than one feed distributor 140 may be employed toprovide feed in various states, e.g., one feed distributor 140 mayprovide feed in a vapor form while a second feed distributor 140 mayprovide feed in a liquid form. Feed distributor 140 includes a body 142from which a plurality of necks 144 extend. Each riser reactor 130 hasat least one associated neck 144. Each head 146 of each neck 144 ispositioned adjacent to the first end 132 of each riser reactor 130.Desirably, each head 146 extends upwardly into each riser reactor 130.More desirably, each head 146 is positioned at or above the mouth 136 atthe first end 132 of each riser reactor 130. Feed distributor 140 mayinclude an optional flow control device, not shown, positioned on feeddistributor 140 to provide an equal amount of feed to each neck 144 or aflow control device may be positioned on each neck 144. The flow controldevice may also be employed to measure flow as well as control. Further,a nozzle, not shown, may be positioned on each head 146 to furthercontrol the distribution of the feed to each riser reactor 130.Additionally, each head 146 may be fitted with a screening device, notshown, to prevent back flow of catalyst into any of necks 144 of feeddistributor 140 and, subsequently, into body 142 of feed distributor140.

[0089] At least one catalyst return 150 provides fluid communicationbetween the separation zone 122 of shell 120 and the riser reactors 130.Each catalyst return 150 has a first end 152 and a second end 154. Thefirst end 152 of the catalyst return 150 opens adjacent the second end126 of shell 120 and the second end 154 of catalyst return 150 opens tothe riser reactors 130. Each catalyst return 150 is provided totransport catalyst from the separation zone 122 of shell 120 to thefirst ends 132 of the riser reactors 130. The apparatus 110 may includeone, two, three, four, five, six or more catalyst returns 150.Typically, although not necessarily, the number of catalyst returns 150corresponds to the number of riser reactors 130. Flow of catalystthrough the catalyst return(s) 150 may optionally be controlled throughthe use of flow control devices, not shown, positioned on each catalystreturn 150. The flow control devices can be any type of flow controldevices currently in use in the art to control catalyst flow throughcatalyst transfer lines. If employed, the flow control device isdesirably a ball valve, a plug valve or a slide valve.

[0090] The apparatus 110 further includes a base 160. In the embodimentshown in FIG. 2, the base 160, the catalyst returns 150 and the firstends 132 of the riser reactors 130 define a catalyst retention zone 162.The second ends 154 of the catalyst returns 150 open to the catalystretention zone 162. The catalyst retention zone 162 is provided toretain catalyst which is used to catalyze the hydrocarbon conversionreaction which is conducted in the apparatus 110. As one of skill in theart will appreciate, the boundary between the catalyst retention zone162 and the catalyst return 150 is fluid and depends, at least in part,on the level of catalyst contained in the catalyst retention zone 162and the catalyst return 150.

[0091] A fluid distributor 170 is also positioned in or near the base160 of the apparatus 110. The fluid distributor 170 includes a conduit172 into which a fluidizing fluid is fed into catalyst retention zone162 to fluidize a fluidizable catalyst contained in the catalystretention zone 162 and the catalyst returns 150. Additional fluiddistributors 170, as shown in FIG. 2, may also be positioned on thecatalyst return(s) 150 to provide additional fluidizing fluid in thecatalyst return(s) 150.

[0092] The HCA 110 may also include an outlet 180 through which thecatalyst can be removed from the apparatus 110. The outlet 180 ispositioned adjacent the quiescent zone 128 in the second end 126 of theshell 120. It is desirable for the outlet 180 to positioned such thatcatalyst can be removed from the shell 120 through the quiescent zone128. The apparatus 110 may also include an inlet 182 through which thecatalyst may be placed into the apparatus 110. Although the inlet 182 isshown as being positioned on the second end 126 of the shell 120, theinlet 182 may be positioned at any position on the apparatus 110. Lines184 are provided to remove products and unreacted feed, if any, from theseparation zone 122 of the HCA 110.

[0093] A series of separation devices 186 are shown as being positionedin the separation zone 122 of shell 120. The separation devices 186 maybe cyclonic separators, filters, screens, impingement devices, plates,cones or any other devices which would separate the catalyst from theproduct of the conversion reaction.

[0094] An impingement device 190 is positioned in the first end 124 ofthe shell 120. The impingement device 190 is provided to direct catalystleaving the riser reactors 130 away from the second ends 134 of theriser reactors 130 and to limit the amount of catalyst falling back intothe riser reactors 130. Desirably, the impingement device 190 ispositioned opposite the second ends 134 of the riser reactors 130.

[0095] A series of supports 192 are also shown in FIG. 2. The supports192 are merely shown to be illustrative of one possible means forsupporting the HCA 110.

[0096] As one of skill in the art will appreciate, the HCA shown in FIG.2 functions similarly to that shown in FIG. I and will not be discussedin detail except to illustrate those features not shown in FIG. 1.

[0097] With reference to FIG. 2, catalyst is provided to the catalystretention zone 162 and is fluidized in the catalyst retention zone 162and the catalyst returns 150 by the fluidizing fluid provided throughthe fluid distributor 170. The feed is provided to the riser reactors130 through the feed distributor 140. The amount of feed provided toeach of the riser reactors 130 is the same as that described above inconjunction with the description of FIG. 1. The catalyst and feed flowupwardly through the riser reactors 130, in the same manner as describedabove in conjunction with the description of the riser reactors 20 inFIG. 1.

[0098] With continuing reference to FIG. 2, the catalyst, product andunreacted feed, if any, exit through the second ends 134 of the riserreactors 130 into the separation zone 122 of the shell 120. At least aportion, and desirably most, of the catalyst contacts the impingementdevice 190 and is deflected toward the sides of the shell 120. Theseparators 186 separate at least a portion of the catalyst from theproduct and unreacted feed. The product and unreacted feed are removedfrom the shell 120 of the hydrocarbon conversion device 10 through thelines 184. The catalyst, which is separated by the separators 186, fallsinto the quiescent zone 128. The remainder of the catalyst is returnedto contact the feed through the catalyst returns 150.

[0099] A portion of the catalyst contained in the quiescent zone 128 canbe removed from one or more of the HCAes and be sent to a catalystregeneration apparatus via outlet 180, such as catalyst regenerationapparatus 90 shown in FIG. 1, or removed from the HCA 110 for furtherprocessing. Additionally, catalyst in the quiescent zone 128 may spillover into the catalyst returns 150 and be returned to contact the feed.

[0100] Another embodiment of one or more of the HCA's of the presentinvention is shown in FIG. 3. The apparatus 200 comprises a shell 212, aplurality of riser reactors 220, feed distributors 230, and a catalystreturn 250.

[0101] With continuing reference to FIG. 3, the shell 212 defines aseparation zone 214 in which a product of the hydrocarbon conversionreaction is separated from the catalyst which catalyzes the hydrocarbonconversion reaction. Shell 212 includes a first end 216 and a second end218.

[0102] Riser reactors 220 extend into shell 212 and the separation zone214. Each riser reactor 220 includes a first end 222 into which thecatalyst and feed are fed to conduct the hydrocarbon conversionreaction. Each riser reactor 220 further includes a second end 224through which the catalyst, product, and unreacted feed, if any, exitthe riser reactor 220. The first end 222 of each riser reactor 220terminates in a mouth 226 through which the catalyst and feed are fedinto the riser reactor 220. The number and dimensions of the riserreactors 220 is discussed above in conjunction with the description ofFIG. 1.

[0103] With continuing reference to FIG. 3, to provide a feed to theriser reactors 220, at least one feed distributor 230 is positioned nearthe first ends 222 of the riser reactors 220. More than one feeddistributor 230 may be employed to provide feed in various states, e.g.,one feed distributor 230 may provide feed in a vapor form while a secondfeed distributor 230 may provide feed in a liquid form. Each feeddistributor includes a body, not shown, from which at least one neck 232extends. Each riser reactor 220 has at least one associated neck 232.Each feed distributor 230 terminates in a head 234. Each head 234 ispositioned adjacent to the first end 222 of each riser reactor 220.Desirably, each head 234 extends upwardly into each riser reactor 220.More desirably, each head 234 is positioned at or above the mouth 226 ofthe first end 222 of each riser reactor 220. Feed distributor 230 mayinclude an optional flow control device, not shown, positioned on feeddistributor 230 to provide an equal amount of feed to each head 234. Theflow control device can also be employed to measure flow as well.Further, a nozzle, not shown, may be positioned on each head 234 tofurther control the distribution of the feed to each riser reactor 220.Additionally, each head 234 may be fitted with screening device, notshown, to prevent back flow of catalyst into any of the feeddistributors 230.

[0104] In the HCA 200 shown in FIG. 3, a single catalyst return 250 ispositioned centrally in relation to the riser reactors 220. The catalystreturn 250 provides fluid communication between the separation zone 214of the shell 212 and the riser reactors 220. The catalyst return 250 hasa first end 252 and a second end 254. The first end 252 of the catalystreturn 250 opens into the first end 214 of shell 212 and the second end254 of catalyst return 250 opens to the riser reactors 220. A series ofarms 256 or standpipes are positioned on the second end 254 of thecatalyst return 250. The arms 256 extend from the catalyst return 250 toeach of the riser reactors 220 and provide fluid communication betweenthe catalyst return 250 and the riser reactors 220. The number of arms256 will correspond to the number of riser reactors 220 with each riserreactor 230 having at least one corresponding arm 256. The catalystreturn 250 is provided to transport catalyst from the separation zone214 of shell 212 to the first ends 222 of the riser reactors 220. Flowof catalyst through the catalyst return 250 may optionally be controlledthrough the use of a flow control device 258 positioned on the catalystreturn 250 or on each arm 256. The flow control device(s) 258 can be anytype of flow control devices currently in use in the art to controlcatalyst flow through catalyst transfer lines. If employed, the flowcontrol device 258 is desirably a ball valve, a plug valve or a slidevalve.

[0105] In the embodiment shown in FIG. 3, the first end 252 of thecatalyst return 250 and the arms 256 define a catalyst retention zone262. The arms 256 of the catalyst return 250 open to the catalystretention zone 262. The catalyst retention zone 262 is provided toretain catalyst which is used to catalyze the hydrocarbon conversionreaction which is conducted in the apparatus 200. As one of skill in theart will appreciate, the boundary between the catalyst retention zone262 and the catalyst return 250 is fluid and depends, at least in part,on the level of catalyst contained in the catalyst retention zone 262and the arms 256 of the catalyst return 250.

[0106] At least one fluid distributor 270 is positioned beneath thecatalyst retention zone 262. The fluid distributor 270 includes aconduit 272 into which a fluidizing fluid is fed to fluidize afluidizable catalyst in the catalyst retention zone 262 and the catalystreturn 250. Additional fluid distributors 270, as shown in FIG. 3, mayalso be positioned on the catalyst return 250 to further fluidizefluidizable catalyst contained in the catalyst return 250.

[0107] The HCA 200 may also include an outlet 280 through which thecatalyst can be removed from the apparatus 200. The outlet 280 is shownas being positioned on the second end 218 of the shell 212 but may bepositioned at any position on the apparatus 200. The apparatus 200 mayalso include an inlet 282 through which the catalyst may be placed intothe apparatus 200. Although the inlet 282 is shown as being positionedon the second end 218 of the shell 212, the inlet 282 may be positionedat any position on the apparatus 200. A line 284 may be provided toremove products from the apparatus 200.

[0108] A series of separation devices 286 are shown as being positionedin the separation zone 214 of shell 212. The separation devices 286 maybe cyclonic separators, filters, screens, impingement devices, plates,cones or any other devices which would separate the catalyst from theproduct of the conversion reaction. The separation devices 286 are shownin FIG. 3 as cyclonic separators 288.

[0109] A series of supports 292 are also shown in FIG. 3. The supports292 are merely shown to be illustrative of one possible means forsupporting the HCA 200.

[0110] The HCA 200 which is shown in FIG. 3 functions similarly to thatshown in FIGS. 1 and 2. The apparatus 200 shown in FIG. 3 functions inthe following manner.

[0111] The apparatus 200 is filled with an appropriate amount ofcatalyst which is retained in the catalyst return 250 and the catalystretention zone 262. The catalyst is fluidized in the catalyst return 250and the catalyst retention zone 262 by means of a fluidizing fluid whichis provided to the HCA 200 through the conduits 272 of the fluiddistributors 270. The flow of catalyst to the riser reactors 220 can becontrolled by the flow control devices 258. Feed is provided to theriser reactors 220 through the feed distributors 230. The amount of feedprovided to the riser reactors 220 is the same as that discussed abovein conjunction with the description of FIG. 1. The feed and the catalystflow upwardly in the riser reactors 230 by the principle of eductionwhich is also described above.

[0112] The catalyst, product and unreacted feed, if any, exit the riserreactors 220 through their second ends 224. The catalyst is separatedfrom the product and any unreacted feed by the separation devices 286.The separated catalyst is fed to the second end 218 of shell 212 whilethe product and any unreacted feed are removed from the apparatusthrough the line 284.

[0113] A portion of the catalyst may be removed from the apparatus 200through the outlet 280 and sent to a regeneration apparatus, not shown,or removed entirely from the apparatus 200. The regenerated catalyst isreturned to the apparatus 200 through the inlet 282.

[0114] The separated catalyst enters the first end 252 of the catalystreturn 250 and is recycled to be reused in the hydrocarbon conversionreaction. The catalyst is returned through the catalyst return 250 tothe catalyst containment area 262 where the catalyst is maintained in afluidized state by the fluidizing fluid provided through the fluiddistributors 270.

[0115] Another embodiment of one or more of the HCA's 300 is shown inFIG. 4. The apparatus 300 comprises a shell 310, a plurality of riserreactors 330, a feed distributor 340 and a fluid distributor 350.

[0116] With continued reference to FIG. 4, the shell 310 is formed by awall 312 and is hollow. Shell 310 has a first end 314 and a second end316. The first end 314 of shell 310 defines a separation zone 318 inwhich the catalyst is separated from the product of the hydrocarbonconversion reaction. The shell 310 further includes a wall extension320, which extends upwardly into the first end 314 of shell 310 from thesecond end 316 of shell 310, and a funnel portion 322. The wallextension 320 and the funnel portion 322 define a quiescent zone 324 inwhich a portion of the catalyst can be retained prior to being removedfrom the shell 310.

[0117] In the embodiment shown in FIG. 4, a plurality of riser reactors330 are positioned inside shell 310, as shown in FIG. 4. Each riserreactor 330 extends substantially parallel to a longitudinal axis ofshell 310 and has a wall 331. Each riser reactor 330 has a first end 332and a second end 334. The first end 332 of each riser reactor 330 ispositioned in the second end 316 of shell 310. The second end 334 ofeach riser reactor 330 extends into the first end 314 of shell 310. Thefirst end 332 of each riser reactor 330 terminates in a mouth 335through which the catalyst and feed are fed into the riser reactor 330.Although the HCA 300 is shown in FIG. 4 as containing three riserreactors 330, apparatus 300 desirably contains two or more riserreactors 330. The number and size of the riser reactors 330 is describedin conjunction with the description of FIG. 1.

[0118] With continuing reference to FIG. 4, wall 312 of shell 310 andwall 331 of each of the riser reactors 330 define a catalyst retentionzone 336. The catalyst retention zone 336 contains the catalyst utilizedto catalyze the hydrocarbon conversion reaction. When the apparatus 300is in operation, catalyst retention zone 336 contains the catalyst in afluidized state, as will be described in detail below. Wall extension320, wall 312 of the shell 310 and the walls 331 of each of the riserreactors 330 also define a catalyst return 338. The catalyst return 338directs catalyst which has been used in a conversion reaction from theseparation zone 318 in the first end 314 of the shell 310 to thecatalyst retention zone 336. As one of skill in the art will appreciate,the boundary between the catalyst retention zone 336 and the catalystreturn 338 is fluid and depends, at least in part, on the level ofcatalyst contained in the catalyst retention zone 336.

[0119] To provide a feed to the riser reactors 330, at least one feeddistributor 340 is positioned near the first ends 332 of the riserreactors 330. More than one feed distributor 340 may be employed toprovide feed in various states, e.g., one feed distributor 340 mayprovide feed in a vapor form while a second feed distributor 340 mayprovide feed in a liquid form. Feed distributor 340 includes a body 342from which a plurality of necks 344 extend. Each riser reactor 330 hasat least one associated neck 344. Each neck 344 terminates in a head346. Each head 346 of each neck 344 is positioned adjacent to the firstend 332 of each riser reactor 330. Desirably, each head 346 extends intoeach riser reactor 330. More desirably, each head 346 is positioned ator above the mouth 335 at the first end 332 of each riser reactor 330.Feed distributor 340 may include an optional flow control device 348positioned on feed distributor 340 to provide an equal amount of feed toeach neck 344 and, if desired, to measure the flow through each neck344. As shown in FIG. 4, the flow control device 348 is a valve 350.Useful types of valves are described above. Further, a nozzle, notshown, may be fitted onto each head 346 to distribute the feed into eachriser reactor 330. Additionally, each head 346 may be fitted withscreening device, not shown, to prevent back flow of catalyst into anyof necks 344 of feed distributor 340 and, subsequently into body 342 offeed distributor 340.

[0120] A fluid distributor 350 is also positioned in second end 316 ofshell 310. The fluid distributor 350 includes a conduit 352 into which afluidizing fluid is fed to fluidize a fluidizable catalyst in thecatalyst retention zone 336 and the catalyst return 338. An optionaldisperser 354 may be positioned between the fluid distributor 350 andthe catalyst retention zone 336 to disperse the fluidizing fluid aboutthe catalyst retention zone 336 and the catalyst return 338. Disperser354 is desirably positioned perpendicular to the longitudinal axis ofshell 310 in the second end 316 of shell 310. Disperser 354 may be ascreen, a grid, a perforated plate or similar device through which thefluidizing fluid is fed to provide even distribution of the fluidizingfluid to the catalyst retention zone 336.

[0121] To separate products from the hydrocarbon conversion reactionfrom the catalyst, a separator 360 or series of separators 360, may bepositioned in first end 314 of shell 310. The separators 360 are shownin FIG. 4 as being cyclonic separators 362. Other types of separators360 such as filters, screens, impingement devices, plates, cones andother devices which would separate the products from the catalyst mayalso be positioned in the first end 314 of shell 310. The number ofseparators 360 depends upon the desired operating efficiency, particlesize of the catalyst, the gas superficial velocity, production capacity,and other parameters. The products are removed from shell 310 through aline 364 or a plurality of lines 364 for further processing such as, forexample, separation and purification.

[0122] The apparatus 300 may further include an outlet 370 through whichcatalyst may be removed from the shell 310 and an inlet 372 throughwhich catalyst may be placed into shell 310. The positioning of outlet370 and inlet 372 is not critical. However, it is desirable for theoutlet 370 to be positioned such that catalyst can be removed from theshell 310 through the quiescent zone 324.

[0123] An impingement device 380 is positioned in the first end 314 ofthe shell 310. The impingement device 380 is provided to direct catalystleaving the riser reactors 330 away from the second ends 334 of theriser reactors 330 and to limit the amount of catalyst falling back intothe riser reactors 330.

[0124] A support 392 is also shown in FIG. 4. The support 392 is merelyshown to be illustrative of one possible means for supporting the HCA300.

[0125] As shown in FIG. 4, one or more of the HCA's includes anassociated catalyst regeneration apparatus 90 which is in fluidcommunication with the HCA 300. The catalyst regeneration apparatus 90includes a catalyst regenerator 92, which is in fluid communication withthe HCA 300 and an optional catalyst stripper 94, which is in fluidcommunication with the catalyst regenerator 92 and which may be in fluidcommunication with the HCA 300. A first line 96 provides fluidcommunication between the catalyst stripper 94 and shell 310 throughoutlet 370. Optionally, the catalyst stripper 94 may include a secondinput line (not shown) from a second HCA (not shown). In this manner,catalyst from the second HCA can be transported by the second input lineto the catalyst stripper 94.A second line 98 provides fluidcommunication between the catalyst stripper 94 and the catalystregenerator 92. A third line 100 provides fluid communication betweenthe catalyst regenerator 92 and the inlet 372 on shell 310. A flowcontrol device 102 may optionally be positioned on first line 96 tocontrol the flow of catalyst between the shell 12 and the catalyststripper 94. A flow control device 104 may optionally be positioned onsecond line 98 to control the flow of catalyst between the catalyststripper 94 and the catalyst regenerator 92.

[0126] Although the catalyst stripper 94 is shown on FIG. 4 as beingseparate from the catalyst regenerator 92, one skilled in the art willappreciate that the catalyst stripper 94 may be integrally formed withthe catalyst regenerator 92. One skilled in the art will also appreciatethat, although FIG. 4 shows third line 100 as returning the catalyst tothe separation zone 318 through line 372, the catalyst may also bereturned to the catalyst return 338, the catalyst retention zone 336 andcombinations of the separation zone 318, the catalyst return 338 and thecatalyst retention zone 336.

[0127] When in operation, one or more of the HCA's, as shown in FIG. 4,functions in the following manner. The catalyst retention zone 336 isfilled with a catalyst suitable to carry out the desired hydrocarbonconversion reaction. The catalyst should be of a type which isfluidizable. To fluidize the catalyst in the catalyst retention zone 336and the catalyst return 338, a fluidizing fluid is fed into the fluiddistributor 350 through conduit 352. The fluidizing fluid is dispersedwithin the shell 310 of the HCA 300 by the disperser 354. Usefulfluidizing fluids include, but are not limited to, nitrogen, steam,carbon dioxide and hydrocarbons. The choice of fluidizing fluid dependsupon the type of conversion reaction being conducted in the hydrocarbonconversion apparatus 300.

[0128] Once the catalyst has reached an acceptable fluidized state, afeed is fed into the HCA 300 through feed distributor 340. The feedenters the body 342 of feed distributor 340, passes through the necks344 of feed distributor 340 and exits through the heads 346 of feeddistributor 340. The feed is distributed to each of the riser reactors330 through the mouths 335 at the first ends 332 of the riser reactors330.

[0129] A pressure differential created by the velocity of the feedentering the first ends 332 of the riser reactors 330 and the pressureof the height of fluidizable catalyst in the catalyst retention zone 336causes catalyst to be aspirated into the first ends 332 of the riserreactors 330. The catalyst is transported through the riser reactors 330under well known principles of eduction in which the kinetic energy ofone fluid, in this case the feed, is used to move another fluid, in thiscase the fluidized catalyst. The catalyst and feed travel from the firstends 332 to the second ends 334 of the riser reactors 330. As thecatalyst and feed travel through the riser reactors 330, the hydrocarbonconversion reaction occurs and a conversion product is produced.

[0130] The conversion product(s), unreacted feed, if any, and thecatalyst exit the riser reactors 330 through their second ends 334 andenter the catalyst separation zone 318 in the first end 314 of shell310. In the catalyst separation zone 318, the conversion product andunreacted feed, if any, are separated from the catalyst by the separator360. Desirably, the conversion product and unreacted feed, if any, areseparated by a series of cyclonic separators 362 as shown in FIG. 4.Further, at least a portion of the catalyst exiting the riser reactors330 contacts the impingement device 380 and is deflected away from thesecond ends 334 of the riser reactors 330 to the quiescent zone 324.

[0131] Once the catalyst has been separated from the conversion productand the unreacted feed, if any, are removed from the shell 310 throughthe lines 364 for further processing such as separation andpurification. A portion of the catalyst falls to the quiescent zone 324in which the catalyst is retained until it is removed from the shell310. The catalyst is removed from the quiescent zone 324 through outlet370 and can be sent for regeneration in the catalyst regenerationapparatus 90.

[0132] The catalyst regenerator 92 preferably receives catalyst from aplurality of HCA's as illustrated in FIG. 6, and as described in moredetail below. As shown in FIG. 4, the catalyst regenerator 92 optionallymay receive catalyst via second line 98 and line 101, which is connectedto a second HCA (not shown). Optionally, line 101 may be directed to asecond stropping unit (not shown) which is in fluid communication withthe second HCA. In this embodiment, catalyst may be delivered to thecatalyst regenerator 92 from the second HCA through line 101. Similarly,catalyst from more than two HCA's may be directed to the catalyststripper 94. Additionally or alternatively, catalyst from more than twoHCA's may be directed to the catalyst regenerator 92. The function ofthe catalyst regeneration apparatus 90 is discussed above in conjunctionwith the description of FIG. 1 and will not be discussed in furtherdetail here. A portion of the catalyst in the quiescent zone 324 willfall out of the quiescent zone 324 into the catalyst return 338 and bereturned to contact the feed.

[0133] Returning to FIG. 4, the remaining portion of the catalyst, afterbeing separated from the products and unreacted feed, falls from thefirst end 314 of shell 310 through the catalyst return 338 to thecatalyst retention zone 336. From the catalyst retention zone 336, thecatalyst is recycled for use in the hydrocarbon conversion reaction.

[0134] Representative embodiments of possible configurations of riserreactors and catalyst returns are shown in cross section in FIG. 5. FIG.5A shows a possible configuration for the riser reactors 20 for the HCA10 shown in FIG. 1. As shown in FIG. 5A, the riser reactors 20 arecontained within a shell 26. If contained within a shell 26, the areabetween the riser reactors and the shell 26 is filled with refractorymaterial 28. Useful refractory materials 28 include sand, cement,ceramic materials, high alumina bricks containing mullite or corundum,high silica bricks, magnesite bricks, insulating firebrick of clay orkaolin or any other high temperature resistant material.

[0135]FIG. 5B shows a cross section of a HCA similar to apparatus 10shown in FIG. 1. In this embodiment, the riser reactors 20 are againcontained within a shell 26. The shell 26 is filled with refractorymaterial 28 as described above. In this embodiment, the catalyst returns50 are also contained within the shell 26 and surrounded by therefractory material 28.

[0136]FIG. 5C shows a possible configuration for the riser reactors 220shown in FIG. 3. In the embodiment shown in FIG. 5C, the catalyst return250 is shown as being centrally positioned in relation to the riserreactors 220. The riser reactors 220 and the catalyst return 250 arecontained within a shell 226. The area between the riser reactors andthe shell 226 is filled with refractory material 228. Useful refractorymaterials are described above in conjunction with the description ofFIG. 5A.

[0137]FIG. 5D shows a possible configuration for the riser reactors 330shown in FIG. 4. As shown in FIG. 5D, the riser reactors 330 arecentrally located within the shell 310. As described above inconjunction with the description of FIG. 4, the walls 331 of the riserreactors 330 and the shell 310 define the catalyst return 338. The areabetween the riser reactors 330 can optionally be filled with a firstrefractory material 382. The shell 310 may also be optionally filledwith a second refractory material 384. Useful refractory materials aredescribed above in conjunction with the description of FIG. 5A. Withcontinuing reference to FIG. 5D, a person of skill in the art willappreciate that the first refractory material 382 and the secondrefractory material 384 can be the same or different material.

[0138]FIG. 5E shows another possible configuration for the riserreactors 330 shown in FIG. 4. As shown in FIG. 5E, the riser reactors330 are centrally located within the shell 310. In this embodiment, theriser reactors 330 are contained within a second shell 386 which has awall 388. The catalyst return 338 is defined by the wall 388 of thesecond shell 386 and the shell 310. The areas between the walls 331 ofthe riser reactors 330 and the wall 388 of the second shell 386 arefilled with a first refractory material 390. The shell 310 may also befilled with a second refractory material 392. Useful refractorymaterials are described above in conjunction with the description ofFIG. 5A. With continuing reference to FIG. 5E, a person of skill in theart will appreciate that the first refractory material 390 and thesecond refractory material 392 can be the same or different material.

[0139] While the riser reactors and catalyst returns are shown in thevarious Figures as having a circular cross section, the riser reactorsand catalyst returns may have any cross section which would facilitateoperation of the HCA. Other useful cross sections for the riser reactorsand the catalyst returns include elliptical cross sections, polygonalcross sections and cross sections of sections of ellipses and polygons.Desirable cross-sections for the riser reactors and catalyst returnsinclude circles and regular polygons with sides of equal lengths. By“regular”, it is meant that the shape of the cross-section has no linesegments with vertices, inside the boundaries of the shape, havingangles greater than 180°. The most desirable cross-sections are circles,and triangles, squares, and hexagons with sides of equal length. Themeans of determining cross-sectional areas for any cross-section shapeis based on long established geometric principles well known to thoseskilled in the art. Similarly, desirable cross-sections for theseparation zone include circles and regular polygons with sides of equallengths. The most desirable cross-sections are circles, and triangles,squares, and hexagons with sides of equal length.

[0140] While the position of the riser reactors relative to theseparation zone are shown in the figures as equidistant and symmetrical,alternate configurations are within the scope of the present invention.For example, the riser reactors may be positioned on one side of theseparation zone in a hemispherical layout. As another example, when theseparation zone has a circular or approximately circular cross-section,the riser reactors may be positioned in a line along the diameter theseparation zone. One skilled in the art will appreciate that a widevariety of configurations of the risers relative to the separation zonemay be utilized in the present invention.

[0141] One skilled in the art will further appreciate that the optionalmultiple riser reactors in a given HCA of the present invention may beformed by dividing a single riser reactor into a plurality of smallerriser reactors. For example, a larger, reactor having a circular crosssection could be divided into several pie-shaped riser reactors. Asanother example, a riser reactor having a square cross section could bedivided into a plurality of riser reactors having either rectangular orsmaller square cross sections.

[0142] The optional multiple riser HCA's of the present invention areuseful to conduct most any hydrocarbon conversion process in which afluidized catalyst is employed. Typical reactions include, for example,olefin interconversion reactions, oxygenate to olefin conversionreactions (e.g., MTO reactions), oxygenate to gasoline conversionreactions, malaeic anhydride formulation, vapor phase methanolsynthesis, phthalic anhydride formulation, Fischer Tropsch reactions,and acrylonitrile formulation. One or more of these hydrocarbonconversion processes may be coupled to, e.g., in fluid communicationwith the integrated catalyst regeneration system in accordance with thepresent invention. For example, in one embodiment, a first HCA produceslight olefins through an MTO reaction process, while a second HCAperforms an olefin interconversion process. Both the first and secondHCA's may send catalyst to an integrated regeneration system.

[0143] The process for converting oxygenates to olefins employs a feedincluding an oxygenate. As used herein, the term “oxygenate” is definedto include, but is not necessarily limited to, hydrocarbons containingoxygen such as the following: aliphatic alcohols, ethers, carbonylcompounds (aldehydes, ketones, carboxylic acids, carbonates, and thelike), and mixtures thereof. The aliphatic moiety desirably shouldcontain in the range of from about 1-10 carbon atoms and more desirablyin the range of from about 1-4 carbon atoms. Representative oxygenatesinclude, but are not necessarily limited to, lower molecular weightstraight chain or branched aliphatic alcohols, and their unsaturatedcounterparts. Examples of suitable oxygenates include, but are notnecessarily limited to the following: methanol; ethanol; n-propanol;isopropanol; C₄-C₁₀ alcohols; methyl ethyl ether; dimethyl ether;diethyl ether; di-isopropyl ether; methyl formate; formaldehyde;di-methyl carbonate; methyl ethyl carbonate; acetone; and mixturesthereof. Desirably, the oxygenate used in the conversion reaction isselected from the group consisting of methanol, dimethyl ether andmixtures thereof. More desirably the oxygenate is methanol. The totalcharge of feed to the riser reactors may contain additional components,such as diluents.

[0144] One or more diluents may be fed to the riser reactors with theoxygenates, such that the total feed mixture comprises diluent in arange of from about 1 mol % and about 99 mol %. Diluents which may beemployed in the process include, but are not necessarily limited to,helium, argon, nitrogen, carbon monoxide, carbon dioxide, hydrogen,water, paraffins, other hydrocarbons (such as methane), aromaticcompounds, and mixtures thereof. Desired diluents include, but are notnecessarily limited to, water and nitrogen.

[0145] A portion of the feed may be provided to the reactor in liquidform. When a portion of the feed is provided in a liquid form, theliquid portion of the feed may be either oxygenate, diluent or a mixtureof both. The liquid portion of the feed may be directly injected intothe individual riser reactors, or entrained or otherwise carried intothe riser reactors with the vapor portion of the feed or a suitablecarrier gas/diluent. By providing a portion of the feed (oxygenateand/or diluent) in the liquid phase, the temperature in the riserreactors can be controlled. The exothermic heat of reaction of oxygenateconversion is partially absorbed by the endothermic heat of vaporizationof the liquid portion of the feed. Controlling the proportion of liquidfeed to vapor feed fed to the reactor is one possible method forcontrolling the temperature in the reactor and in particular in theriser reactors.

[0146] The amount of feed provided in a liquid form, whether fedseparately or jointly with the vapor feed, is from about 0.1 wt. % toabout 85 wt. % of the total oxygenate content plus diluent in the feed.More desirably, the range is from about 1 wt. % to about 75 wt. % of thetotal oxygenate plus diluent feed, and most desirably the range is fromabout 5 wt. % to about 65 wt. %. The liquid and vapor portions of thefeed may be the same composition, or may contain varying proportions ofthe same or different oxygenates and same or different diluents. Oneparticularly effective liquid diluent is water, due to its relativelyhigh heat of vaporization, which allows for a high impact on the reactortemperature differential with a relatively small rate. Other usefuldiluents are described above. Proper selection of the temperature andpressure of any appropriate oxygenate and/or diluent being fed to thereactor will ensure at least a portion is in the liquid phase as itenters the reactor and/or comes into contact with the catalyst or avapor portion of the feed and/or diluent.

[0147] Optionally, the liquid fraction of the feed may be split intoportions and introduced to riser reactors a multiplicity of locationsalong the length of the riser reactors. This may be done with either theoxygenate feed, the diluent or both. Typically, this is done with thediluent portion of the feed. Another option is to provide a nozzle whichintroduces the total liquid fraction of the feed to the riser reactorsin a manner such that the nozzle forms liquid droplets of an appropriatesize distribution which, when entrained with the gas and solidsintroduced to the riser reactors, vaporize gradually along the length ofthe riser, reactors. Either of these arrangements or a combinationthereof may be used to better control the temperature differential inthe riser reactors. The means of introducing a multiplicity of liquidfeed points in a reactor or designing a liquid feed nozzle to controldroplet size distribution is well known in the art and is not discussedhere.

[0148] The catalyst suitable for catalyzing an oxygenate-to-olefinconversion reaction includes a molecular sieve and mixtures of molecularsieves. Molecular sieves can be zeolitic (zeolites) or non-zeolitic(non-zeolites). Useful catalysts may also be formed from mixtures ofzeolitic and non-zeolitic molecular sieves. Desirably, the catalystincludes a non-zeolitic molecular sieve. Desired molecular sieves foruse with an oxygenate to olefins conversion reaction include “small” and“medium” pore molecular sieves. “Small pore” molecular sieves aredefined as molecular sieves with pores having a diameter of less thanabout 5.0 Angstroms. “Medium pore” molecular sieves are defined asmolecular sieves with pores having a diameter from about 5.0 to about10.0 Angstroms.

[0149] Useful zeolitic molecular sieves include, but are not limited to,mordenite, chabazite, erionite, ZSM-5, ZSM-34, ZSM-48 and mixturesthereof. Methods of making these molecular sieves are known in the artand need not be discussed here. Structural types of small pore molecularsieves that are suitable for use in this invention include AEI, AFT,APC, ATN, ATT, ATV, AWW, BIK, CAS, CHA, CHI, DAC, DDR, EDI, FRI, GOO,KFI, LEV, LOV, LTA, MON, PAU, PHI, RHO, ROG, THO, and substituted formsthereof. Structural types of medium pore molecular sieves that aresuitable for use in this invention include MFI, MEL, MTW, EUO, MTT, HEU,FER, AFO, AEL, TON, and substituted forms thereof.

[0150] Silicoaluminophosphates (“SAPOs”) are one group of non-zeoliticmolecular sieves that are useful in an oxygenate to olefins conversionreaction. SAPOs comprise a three-dimensional microporous crystalframework structure of [SiO₂], [AlO₂] and [PO₂] tetrahedral units. Theway Si is incorporated into the structure can be determined by ²⁹Si MASNMR. See Blackwell and Patton, J. Phys. Chem., 92, 3965 (1988). Thedesired SAPO molecular sieves will exhibit one or more peaks in the ²⁹SiMAS NMR, with a chemical shift [(Si)] in the range of −88 to −96 ppm andwith a combined peak area in that range of at least 20% of the totalpeak area of all peaks with a chemical shift [(Si)] in the range of −88ppm to −115 ppm, where the [(Si)] chemical shifts refer to externaltetramethylsilane (TMS).

[0151] It is desired that the silicoaluminophosphate molecular sieveused in such a process have a relatively low Si/Al₂ ratio. In general,the lower the Si/Al₂ ratio, the lower the C₁-C₄ saturates selectivity,particularly propane selectivity. A Si/Al₂ ratio of less than 0.65 isdesirable, with a Si/Al₂ ratio of not greater than 0.40 being preferred,and a SiAl₂ ratio of not greater than 0.32 being particularly preferred.

[0152] Silicoaluminophosphate molecular sieves are generally classifiedas being microporous materials having 8, 10, or 12 membered ringstructures. These ring structures can have an average pore size rangingfrom about 3.5-15 angstroms. Preferred are the small pore SAPO molecularsieves having an average pore size ranging from about 3.5 to 5angstroms, more preferably from 4.0 to 5.0 angstroms. These pore sizesare typical of molecular sieves having 8 membered rings.

[0153] In general, silicoaluminophosphate molecular sieves comprise amolecular framework of corner-sharing [SiO₂], [AlO₂], and [PO₂]tetrahedral units. This type of framework is effective in convertingvarious oxygenates into olefin products.

[0154] Suitable silicoaluminophosphate molecular sieves for use in anoxygenate to olefin conversion process include SAPO-5, SAPO-8, SAPO-11,SAPO-16, SAPO-17, SAPO-18, SAPO-20, SAPO-31, SAPO-34, SAPO-35, SAPO-36,SAPO-37, SAPO-40, SAPO-41, SAPO-42, SAPO-44, SAPO-47, SAPO-56, the metalcontaining forms thereof, and mixtures thereof. Preferred are SAPO-18,SAPO-34, SAPO-35, SAPO-44, and SAPO-47, particularly SAPO-18 andSAPO-34, including the metal containing forms thereof, and mixturesthereof. As used herein, the term mixture is synonymous with combinationand is considered a composition of matter having two or more componentsin varying proportions, regardless of their physical state.

[0155] Additional olefin-forming molecular sieve materials can be mixedwith the silicoaluminophosphate catalyst if desired. Several types ofmolecular sieves exist, each of which exhibit different properties.Structural types of small pore molecular sieves that are suitable foruse in this invention include AEI, AFT, APC, ATN, ATT, ATV, AWW, BIK,CAS, CHA, CHI, DAC, DDR, EDI, ERI, GOO, KFI, LEV, LOV, LTA, MON, PAU,PHI, RHO, ROG, THO, and substituted forms thereof. Structural types ofmedium pore molecular sieves that are suitable for use in this inventioninclude MFI, MEL, MTW, EUO, MTT, HEU, FER, AFO, AEL, TON, andsubstituted forms thereof. Preferred molecular sieves which can becombined with a silicoaluminophosphate catalyst include ZSM-5, ZSM-34,erionite, and chabazite.

[0156] Substituted SAPOs form a class of molecular sieves known as“MeAPSOs,” which are also useful in the present invention. Processes formaking MeAPSOs are known in the art. SAPOs with substituents, such asMeAPSOs, also may be suitable for use in the present invention. Suitablesubstituents, “Me,” include, but are not necessarily limited to, nickel,cobalt, manganese, zinc, titanium, strontium, magnesium, barium, andcalcium. The substituents may be incorporated during synthesis of theMeAPSOs. Alternately, the substituents may be incorporated aftersynthesis of SAPOs or MeAPSOs using many methods. These methods include,but are not necessarily limited to, ion-exchange, incipient wetness, drymixing, wet mixing, mechanical mixing, and combinations thereof.

[0157] Desired MeAPSOs are small pore MeAPSOs having pore size smallerthan about 5 Angstroms. Small pore MeAPSOs include, but are notnecessarily limited to, NiSAPO-34, CoSAPO-34, NiSAPO-17, CoSAPO-17, andmixtures thereof.

[0158] Aluminophosphates (ALPOs) with substituents, also known as“MeAPOs,” are another group of molecular sieves that may be suitable foruse in an oxygenate to olefin conversion reaction, with desired MeAPOsbeing small pore MeAPOs. Processes for making MeAPOs are known in theart. Suitable substituents include, but are not necessarily limited to,nickel, cobalt, manganese, zinc, titanium, strontium, magnesium, barium,and calcium. The substituents may be incorporated during synthesis ofthe MeAPOs. Alternately, the substituents may be incorporated aftersynthesis of ALPOs or MeAPOs using many methods. The methods include,but are not necessarily limited to, ion-exchange, incipient wetness, drymixing, wet mixing, mechanical mixing, and combinations thereof

[0159] The molecular sieve may also be incorporated into a solidcomposition, preferably solid particles, in which the molecular sieve ispresent in an amount effective to catalyze the desired conversionreaction. The solid particles may include a catalytically effectiveamount of the molecular sieve and matrix material, preferably at leastone of a filler material and a binder material, to provide a desiredproperty or properties, e.g., desired catalyst dilution, mechanicalstrength and the like, to the solid composition. Such matrix materialsare often to some extent porous in nature and often have somenonselective catalytic activity to promote the formation of undesiredproducts and may or may not be effective to promote the desired chemicalconversion. Such matrix, e.g., filler and binder, materials include, forexample, synthetic and naturally occurring substances, metal oxides,clays, silicas, aluminas, silica-aluminas, silica-magnesias,silica-zirconias, silica-thorias, silica-beryllias, silica-titanias,silica-alumina-thorias, silica-aluminazirconias, and mixtures of thesematerials.

[0160] The solid catalyst composition preferably comprises about 1% toabout 99%, more preferably about 5% to about 90%, and still morepreferably about 10% to about 80%, by weight of molecular sieve; and anamount of about 1% to about 99%, more preferably about 5% to about 90%,and still more preferably about 10% to about 80%, by weight of matrixmaterial.

[0161] The preparation of solid catalyst compositions, e.g., solidparticles, comprising the molecular sieve and matrix material, isconventional and well known in the art and, therefore, is not discussedin detail here.

[0162] The catalyst may further contain binders, fillers, or othermaterial to provide better catalytic performance, attrition resistance,regenerability, and other desired properties. Desirably, the catalyst isfluidizable under the reaction conditions. The catalyst should haveparticle sizes of from about 5μ to about 3,000μ, desirably from about10μ to about 200μ, and more desirably from about 20μ to about 150μ. Thecatalyst may be subjected to a variety of treatments to achieve thedesired physical and chemical characteristics. Such treatments include,but are not necessarily limited to, calcination, ball milling, milling,grinding, spray drying, hydrothermal treatment, acid treatment, basetreatment, and combinations thereof.

[0163] Desirably, in an oxygenate to olefin conversion reactionconducted in one or more HCA's of the present invention employs a gassuperficial velocity in the riser reactors of greater than 1 meter persecond (m/s). As used herein and in the claims, the term, “gassuperficial velocity,” is defined as the volumetric flow rate ofvaporized feedstock, and any diluent, divided by the reactorcross-sectional area. Because the oxygenate is converted to a productincluding a light olefin while flowing through the reactor, the gassuperficial velocity may vary at different locations within the reactordepending on the total number of moles of gas present and the crosssection of a particular location in the reactor, temperature, pressure,and other relevant reaction parameters. The gas superficial velocity,including any diluents present in the feedstock, is maintained at a rategreater than 1 meter per second (m/s) at any point in the reactor.Desirably, the gas superficial velocity is greater than about 2 m/s.More desirably, the gas superficial velocity is greater than about 2.5m/s. Even more desirably, the gas superficial velocity is greater thanabout 4 m/s. Most desirably, the gas superficial velocity is greaterthan about 8 m/s.

[0164] Maintaining the gas superficial velocity at these rates increasesthe approach to plug flow behavior of the gases flowing in the riserreactors. As the gas superficial velocity increases above 1 m/s, areduction in axial diffusion or back mixing of the gases results from areduction in internal recirculation of solids, which carry gas withthem. (Ideal plug flow behavior occurs when elements of the homogeneousfluid reactant move through a reactor as plugs moving parallel to thereactor axis). Minimizing the back mixing of the gases in the reactorincreases the selectivity to the desired light olefins in the oxygenateconversion reaction.

[0165] When the gas superficial velocity approaches 1 m/s or higher, asubstantial portion of the catalyst in the reactor may be entrained withthe gas exiting the riser reactors. At least a portion of the catalystexiting the riser reactors is recirculated to recontact the feed throughthe catalyst return.

[0166] Desirably, the rate of catalyst, comprising molecular sieve andany other materials such as binders, fillers, etc., recirculated torecontact the feed is from about 1 to about 100 times, more desirablyfrom about 10 to about 80 times, and most desirably from about 10 toabout 50 times the total feed rate, by weight, of oxygenates to thereactor.

[0167] The temperature useful to convert oxygenates to light olefinsvaries over a wide range depending, at least in part, on the catalyst,the fraction of regenerated catalyst in a catalyst mixture, and theconfiguration of the reactor apparatus and the reactor. Although theseprocesses are not limited to a particular temperature, best results areobtained if the process is conducted at a temperature from about 200° C.to about 1000° C., more preferably from about 200° C. to about 700° C.,desirably from about 250° C. to about 600° C., and most desirably fromabout 300° C. to about 500° C. Lower temperatures generally result inlower rates of reaction, and the formation rate of the desired lightolefin products may become markedly slower. However, at temperaturesgreater than 700° C., the process may not form an optimum amount oflight olefin products, and the rate at which coke and light saturatesform on the catalyst may become too high.

[0168] Light olefins will form—although not necessarily in optimumamounts—at a wide range of pressures including, but not limited to,pressures from about 0.1 kPa to about 5 MPa. A desired pressure is fromabout 5 kPa to about 1 MPa and most desirably from about 20 kPa to about500 kPa. The foregoing pressures do not include that of a diluent, ifany, and refer to the partial pressure of the feed as it relates tooxygenate compounds and/or mixtures thereof. Pressures outside of thestated ranges may be used and are not excluded from the scope of theinvention. Lower and upper extremes of pressure may adversely affectselectivity, conversion, coking rate, and/or reaction rate; however,light olefins will still form and, for that reason, these extremes ofpressure are considered part of the present invention.

[0169] A wide range of WHSV's for the oxygenate conversion reaction,defined as weight of total oxygenate fed to the riser reactors per hourper weight of molecular sieve in the catalyst in the riser reactors,function with the present invention. The total oxygenate fed to theriser reactors includes all oxygenate in both the vapor and liquidphase. Although the catalyst may contain other materials which act asinerts, fillers or binders, the WHSV is calculated using only the weightof molecular sieve in the catalyst in the riser reactors. The WHSV isdesirably high enough to maintain the catalyst in a fluidized stateunder the reaction conditions and within the reactor configuration anddesign. Generally, the WHSV is from about 1 hr⁻¹ to about 5000 hr⁻¹,desirably from about 2 hr⁻¹ to about 3000 hr⁻¹, and most desirably fromabout 5 hr⁻¹ to about 1500 hr⁻¹. The applicants have discovered thatoperation of the oxygenate to olefin conversion reaction at a WHSVgreater than 20 hr⁻¹ reduces the methane content in the product slate ofthe conversion reaction. Thus, the conversion reaction is desirablyoperated at a WHSV of at least about 20 hr⁻¹. For a feed comprisingmethanol, dimethyl ether, or mixtures thereof, the WHSV is desirably atleast about 20 hr⁻¹ and more desirably from about 20 hr⁻¹ to about 300hr⁻¹.

[0170] The method of making the preferred olefin product in thisinvention can include the additional step of making the oxygenatecompositions from hydrocarbons such as oil, coal, tar sand, shale,biomass and natural gas. Methods for making the compositions are knownin the art. These methods include fermentation to alcohol or ether,making synthesis gas, then converting the synthesis gas to alcohol orether. Synthesis gas can be produced by known processes such as steamreforming, autothermal reforming and partial oxidization.

[0171] One skilled in the art will also appreciate that the olefinsproduced by the oxygenate-to-olefin conversion reaction of the presentinvention can be polymerized to form polyolefins, particularlypolyethylene and polypropylene. Processes for forming polyolefins fromolefins are known in the art. Catalytic processes are preferred.Particularly preferred are metallocene, Ziegler/Natta and acid catalyticsystems. See, for example, U.S. Pat. Nos. 3,258,455; 3,305,538;3,364,190; 5,892,079; 4,659,685; 4,076,698; 3,645,992; 4,302,565; and4,243,691, the catalyst and process descriptions of each being expresslyincorporated herein by reference. In general, these methods involvecontacting the olefin product with a polyolefin-forming catalyst at apressure and temperature effective to form the polyolefin product.

[0172] A preferred polyolefin-forming catalyst is a metallocenecatalyst. The preferred temperature range of operation is between 50° C.and 240° C. and the reaction can be carried out at low, medium or highpressure, being anywhere from 1 bar to 200 bars. For processes carriedout in solution, an inert diluent can be used, and the preferredoperating pressure range is between 10 and 150 bars, with a preferredtemperature between 120° C. and 230° C. For gas phase processes, it ispreferred that the temperature generally be from 60° C. to 160° C., andthat the operating pressure be from 5 bars to 50 bars.

[0173] In addition to polyolefins, numerous other olefin derivatives maybe formed from the olefins produced by the process of the presentinvention or olefins recovered therefrom. These include, but are notlimited to, aldehydes, alcohols, acetic acid, linear alpha olefins,vinyl acetate, ethylene dichloride and vinyl chloride, ethylbenzene,ethylene oxide, ethylene glycol, cumene, isopropyl alcohol, acrolein,allyl chloride, propylene oxide, acrylic acid, ethylene-propylenerubbers, and acrylonitrile, and trimers and dimers of ethylene,propylene or butylenes. The methods of manufacturing these derivativesare well known in the art, and therefore are not discussed here.

[0174] Persons of ordinary skill in the art will recognize that manymodifications may be made to the present invention without departingfrom the spirit and scope of the present invention. The embodimentsdescribed herein are meant to be illustrative only and should not betaken as limiting the invention, which is defined by the followingclaims.

1. A reactor system, comprising: a plurality of reactor units; aregenerator for converting an at least partially deactivated catalyst toa regenerated catalyst; a first conduit system for transferring the atleast partially deactivated catalyst from the reactor units to theregenerator; and a second conduit system for transferring theregenerated catalyst from the regenerator to the plurality of reactorunits.
 2. The system of claim 1, wherein the first conduit systemincludes a first stripping unit for stripping the at least partiallydeactivated catalyst with a first stripping medium.
 3. The system ofclaim 2, wherein the first conduit system includes a second strippingunit for stripping the at least partially deactivated catalyst with asecond stripping medium.
 4. The system of claim 3, wherein the first andsecond stripping units strip at least partially deactivated catalystsfrom separate reactor units.
 5. The system of claim 1, wherein at leastone of the reactor units includes two riser reactors.
 6. The system ofclaim 1, wherein at least one of the reactor units includes three riserreactors.
 7. The system of claim 1, wherein at least one of the reactorunits includes four riser reactors.
 8. The system of claim 1, wherein atleast one of the reactor units includes five riser reactors.
 9. Thesystem of claim 1, wherein at least one of the reactor units includessix riser reactors.
 10. The system of claim 1, wherein at least one ofthe reactor units includes more than six riser reactors.
 11. The systemof claim 1, wherein the system includes two reactor units.
 12. Thesystem of claim 1, wherein the system includes three reactor units. 13.The system of claim 1, wherein the system includes four reactor units.14. The system of claim 1, wherein the system includes more than fourreactor units.
 15. The system of claim 1, wherein at least one of thereactor units includes a plurality of riser reactors and a catalystretention zone provided to contain catalyst which can be fed to theplurality of riser reactors.
 16. The system of claim 15, wherein eachriser reactor in the at least one of the reactor units includes a firstend into which the catalyst can be fed and a second end through whichthe catalyst can exit the riser reactor, and wherein the at least one ofthe reactor units includes a separation zone into which the second endsof the riser reactors discharge the catalyst and products of a reactionconducted in the at least one of the reactor units, the separation zonebeing provided to separate the catalyst from the products.
 17. Thesystem of claim 16, wherein the at least one of the reactor unitsincludes a catalyst return in fluid communication with the separationzone thereof and the catalyst retention zone thereof.
 18. The system ofclaim 17, wherein the at least one of the reactor units includes a feeddistributor including at least one feed head positioned adjacent to eachof the first ends of the plurality of riser reactors therein.
 19. Thesystem of claim 18, wherein the plurality of riser reactors in the atleast one of the reactor units is contained within a common shell havinga wall.
 20. The system of claim 19, wherein the plurality of riserreactors in the at least one of the reactor units, and the respectivewall, define the catalyst retention zone.
 21. The system of claim 20,wherein the shell of the at least one of the reactor units defines theseparation zone.
 22. The system of claim 20, wherein the wall of theshell of the at least one of the reactor units, and the plurality ofriser reactors therein, define the catalyst return.
 23. The system ofclaim 18, wherein the feed distributor of the at least one of thereactor units provides feed to each of the plurality of riser reactorstherein in substantially equal streams through the at least one feedhead.
 24. The system of claim 18, wherein the feed distributor in the atleast one of the reactor units includes a flow control device whichprovides the feed to each of the plurality of riser reactors thereinthrough the feed heads.
 25. The system of claim 18, wherein the at leastone of the reactor units further includes a fluid distributor in fluidcommunication with the catalyst retention zone thereof, the fluiddistributor being provided to feed a fluidizing fluid to the catalystretention zone to fluidize catalyst contained in the catalyst retentionzone.
 26. The system of claim 25, wherein the at least one of thereactor units further includes a disperser, positioned in the first endof the shell, the disperser being provided to disperse the fluidizingfluid in the catalyst retention zone to fluidize the catalyst.
 27. Thesystem of claim 26, wherein the disperser is a device selected from thegroup consisting of a grid, a screen and a perforated plate.
 28. Thesystem of claim 18, wherein the catalyst return is positioned externallyto the plurality of riser reactors in the at least one of the reactorunits.
 29. The system of claim 28, wherein the number of the catalystreturns in the at least one of the reactor units equals the number ofthe plurality of riser reactors in the at least one of the reactorunits.
 30. The system of claim 18, wherein the at least one of thereactor units includes a plurality of catalyst returns.
 31. The systemof claim 30, wherein the at least one of the reactor units includesthree catalyst returns.
 32. The system of claim 30, wherein the at leastone of the reactor units includes four catalyst returns.
 33. The systemof claim 30, wherein the at least one of the reactor units includes aflow control device positioned on at least one of the catalyst returnsthereof.
 34. The system of claim 30, wherein the at least one of thereactor units further includes a flow control device positioned on eachof the plurality of catalyst returns thereof.
 35. The system of claim18, wherein each of the plurality of riser reactors in the at least oneof the reactor units is contained within a common shell.
 36. The systemof claim 18, wherein the at least one of the reactor units furtherincludes an impingement device positioned in the separation zone, theimpingement device being provided to move catalyst away from the secondends of the plurality of riser reactors thereof to the catalyst return.37. The system of claim 35, wherein the at least one of the reactorunits further includes an impingement device positioned in theseparation zone, the impingement device being provided to move catalystaway from the second ends of the plurality of riser reactors thereof tothe catalyst return.
 38. The system of claim 18, wherein the separationzone further includes a quiescent zone in which catalyst can be retaineduntil the catalyst moves from the separation zone.
 39. The system ofclaim 20, wherein the wall of the shell of the at least one of thereactor units and the plurality of riser reactors therein define aquiescent zone in which catalyst is contained until the catalyst movesfrom the separation zone.
 40. The system of claim 16, wherein the atleast one of the reactor units further includes at least one separatorpositioned in the separation zone.
 41. The system of claim 40, whereinthe separator is selected from the group consisting of a cyclonicseparator, a filter, an impingement device and combinations thereof. 42.The system of claim 1, wherein each of the plurality of riser reactorshas a cross sectional area of no greater than 12 m².
 43. The system ofclaim 42, wherein at least one of the reactor units includes a pluralityof riser reactors, each riser reactor having a cross sectional area ofno greater than 7 m².
 44. The system of claim 43, wherein each of theplurality of riser reactors has a cross sectional area or no greaterthan 3.5 m².
 45. The system of claim 1, wherein at least one of thereactor units includes a plurality of riser reactors, each riser reactorhaving a height of from 10 meters to 70 meters.
 46. The system of claim1, wherein at least one of the reactor units includes a plurality ofriser reactors, each riser reactor having a width of from 1 meter to 3meters.
 47. The system of claim 1, wherein at least one of the reactorunits includes a plurality of riser reactors, each riser reactor havinga cross sectional area and the cross sectional area of one riser reactorvaries by no more than 20% from the cross sectional area of anotherriser reactor in a single reactor unit.
 48. The system of claim 1,wherein at least one of the reactor units includes a plurality of riserreactors, each riser reactor having a cross sectional area and the crosssectional area of one riser reactor varies by no more than 10% from thecross sectional area of another riser reactor in a single reactor unit.49. The system of claim 1, wherein at least one of the reactor unitsincludes a plurality of riser reactors, each riser reactor having across sectional area and the cross sectional area of one riser reactorvaries by no more than 1% from the cross sectional area of another riserreactor in a single reactor unit.
 50. A reactor system, comprising: afirst reaction unit comprising a first plurality of riser reactors; asecond reaction unit comprising a second plurality of riser reactors,wherein each of the first and second reaction units has a first end intowhich a catalyst can be fed and a second end through which the catalystcan exit the reaction unit; a regeneration unit having a regenerationinlet and a regeneration outlet; a regeneration line having a pluralityof first line ends in fluid communication with the second ends of thefirst and second reaction units and a second line end extending to theregeneration inlet; and a return line having a first return end in fluidcommunication with the regeneration outlet, a second return enddirecting a first portion of the catalyst to the first reaction unit,and a third return end directing a second portion of the catalyst to thesecond reaction unit.
 51. The system of claim 50, further comprising: afirst stripping unit having a first stripping inlet in fluidcommunication with the second end of the first reaction unit and a firststripping outlet in fluid communication with the regenerator inlet. 52.The system of claim 51, wherein the first stripping inlet is in fluidcommunication with the second end of the second reaction unit.
 53. Thesystem of claim 52, further comprising: a first stripping return linehaving a first stripping return end in fluid communication with thefirst stripping outlet, and a second stripping return end in fluidcommunication with the regeneration inlet.
 54. The system of claim 51,further comprising: a second stripping unit having a second strippinginlet in fluid communication with the second end of the second reactionunit and a second stripping outlet in fluid communication with theregenerator inlet.
 55. A method for forming olefins in a methanol toolefin reactor system, comprising: contacting in a first reaction unit afirst oxygenate with a first catalyst under conditions effective toconvert at least a portion of the first oxygenate to a first olefin andat least partially deactivating the first catalyst to form a deactivatedfirst catalyst; contacting in a second reaction unit a second oxygenatewith a second catalyst under conditions effective to convert at least aportion of the second oxygenate to a second olefin and at leastpartially deactivating the second catalyst to form a deactivated secondcatalyst; directing the deactivated first catalyst and deactivatedsecond catalyst to a regeneration unit; regenerating the deactivatedfirst catalyst and the deactivated second catalyst to form regeneratedcatalysts; directing a first portion of the regenerated catalysts to thefirst reaction unit; and directing a second portion of the regeneratedcatalysts to the second reaction unit.
 56. The method of claim 55,further comprising: contacting the deactivated first catalyst with afirst stripping medium in a first stripping unit under conditionseffective to remove interstitial hydrocarbons from the deactivated firstcatalyst.
 57. The method of claim 56, further comprising: contacting thedeactivated second catalyst with a second stripping medium in a secondstripping unit under conditions effective to remove interstitialhydrocarbons from the deactivated second catalyst.
 58. The method ofclaim 56, further comprising: contacting the deactivated second catalystwith the first stripping medium in the first stripping unit underconditions effective to remove interstitial hydrocarbons from thedeactivated second catalyst.
 59. The method of claim 56, wherein thefirst stripping medium is selected from the group consisting of steam,nitrogen, helium, argon, methane, CO₂, CO, hydrogen, and mixturesthereof.
 60. The method of claim 57, wherein the first stripping mediumis selected from the group consisting of steam, nitrogen, helium, argon,methane, CO₂, CO, hydrogen, and mixtures thereof.
 61. The method ofclaim 58, wherein the first stripping medium is selected from the groupconsisting of steam, nitrogen, helium, argon, methane, CO₂, CO,hydrogen, and mixtures thereof.
 62. The method of claim 55, wherein thecontacting in the first reaction unit occurs in a plurality of riserreactors.
 63. The method of claim 62, wherein the contacting in thesecond reaction unit occurs in a plurality of riser reactors.
 64. Ahydrocarbon conversion system, comprising: first and second pluralitiesof riser reactors, each of the riser reactors having a first end intowhich a catalyst can be fed and a second end through which the catalystcan exit the riser reactor; first and second catalyst retention zonesprovided to contain catalyst which can be fed to the first and secondplurality of riser reactors, respectively; first and second separationzones into which the second ends of the first and second pluralities ofriser reactors extend, respectively, the separation zones being providedto separate the catalyst from products of a reaction conducted in thefirst and second pluralities of riser reactors; first and secondcatalyst returns in fluid communication with the first and secondseparation zones, respectively, and the first and second catalystretention zones, respectively; a regenerator for regenerating thecatalyst; first and second catalyst outlet lines, each of the outletlines having a first end into which a catalyst can be fed from the firstand second pluralities of riser reactors, respectively, and a second endthrough which the catalyst can enter the regenerator; and first andsecond catalyst return lines, each of the return lines having a firstend into which a catalyst can be fed from the regenerator and a secondend through which the catalyst can enter the first and secondpluralities of riser reactors, respectively.
 65. The system of claim 64,wherein the first catalyst outlet line includes a first stripping unitfor stripping the catalyst with a first stripping medium.
 66. The systemof claim 65, wherein the second catalyst outlet line includes a secondstripping unit for stripping the catalyst with a second strippingmedium.
 67. The system of claim 65, wherein the second catalyst outletline includes the first stripping unit for stripping the catalyst withthe first stripping medium.
 68. A catalyst regenerator system,comprising: a regeneration zone for contacting an at least partiallydeactivated catalyst with a regeneration medium under conditionseffective to form a regenerated catalyst; a plurality of catalyst inletsfor receiving the at least partially deactivated catalyst from aplurality of reactor units; and a plurality of catalyst outlets fordelivering the regenerated catalyst to the plurality of reactor units.69. The catalyst regenerator system of claim 68, wherein at least one ofthe reactor units comprises a plurality of riser reactors.
 70. Thecatalyst regenerator system of claim 69, wherein two or more of thereactor units comprise a plurality of riser reactors.
 71. The catalystregenerator system of claim 70, wherein three or more of the reactorunits comprise a plurality of riser reactors.
 72. The catalystregenerator system of claim 71, wherein four or more of the reactorunits comprise a plurality of riser reactors.
 73. The catalystregenerator system of claim 68, further comprising: a stripping zone forcontacting the at least partially deactivated catalyst with a strippingmedium under conditions effective to remove interstitial hydrocarbonsfrom the deactivated catalyst.
 74. The catalyst regenerator system ofclaim 68, further comprising: a plurality of stripping zones forcontacting the at least partially deactivated catalyst with a strippingmedium under conditions effective to remove interstitial hydrocarbonsfrom the deactivated catalyst.
 75. A method for regenerating catalyst,comprising: receiving an at least partially deactivated catalyst from aplurality of multiple riser reaction units; heating the catalyst underconditions effective to convert the at least partially deactivatedcatalyst to a regenerated catalyst; and directing the regeneratedcatalyst to the plurality of multiple riser reaction units.
 76. Themethod of claim 75, further comprising: contacting the at leastpartially deactivated catalyst with a stripping medium under conditionseffective to remove interstitial hydrocarbons from the deactivatedcatalyst.
 77. A hydrocarbon conversion system, comprising: a pluralityof reaction units, each unit comprising a plurality of riser reactors;and at least one regeneration unit coupled to the reaction units;wherein the number of reaction units is greater than the number ofregeneration units.
 78. The hydrocarbon conversion system of claim 77,further comprising: at least one stripping unit coupled to the reactionunits.
 79. The hydrocarbon conversion system of claim 78, wherein thenumber of reaction units is greater than the number of stripping units.